Economic, energy, and environmental impacts of alcohol dehydration technology on biofuel production from brown algae

Energy 93 (2015) 2321e2336

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Economic, energy, and environmental impacts of alcohol dehydration technology on biofuel production from brown algae Peyman Fasahati, J. Jay Liu* Department of Chemical Engineering, Pukyong National University, 365 Sinseon-ro, Nam-gu, Busan 608-739, Republic of Korea

a r t i c l e i n f o

a b s t r a c t

Article history: Received 14 January 2015 Received in revised form 28 June 2015 Accepted 24 October 2015 Available online xxx

This study evaluates the impact of alcohol recovery technology on the economics, energy consumption, and environment of bioethanol production from brown algae. The process under consideration is the anaerobic digestion of brown algae to produce VFAs (volatile fatty acids), which are then hydrogenated to produce mixed alcohols. Three alternative processes, i.e., hybrid pervaporation/distillation (PV), hybrid vapor-permeation/distillation (VP), and classical molecular-sieves/distillation (classical), are considered for the dehydration and recovery of ethanol. The alternatives are analyzed in terms of product value (i.e., minimum ethanol selling price e MESP), capital costs, energy consumption, and carbon footprint. For a plant scale of 400,000 ton/year of dry brown algae, the MESPs for the PV (Pervaporation), VP (vapor permeation), and classical processes were calculated to be $1.06/gal, $1.08/gal, and $1.24/gal, respectively. Results show that the PV, VP, and classical processes have $2.0, $2.6, and $4.6 million/year utility costs, respectively, for the recovery of alcohols and produce 23.1, 30.2, and 62.2 kton CO2-eq/year greenhouse gases. Therefore, PV is more economical and environmentally friendly process, with lower MESP, CO2 emissions, and utility requirements. A sensitivity analysis indicates that the selling price of the heavier alcohols and biomass price have the highest impact on the economics of bioethanol production from brown algae. © 2015 Elsevier Ltd. All rights reserved.

Keywords: Bioethanol Brown algae Pervaporation Vapor permeation Techno-economic analysis Carbon footprint

1. Introduction The global energy demand is growing continually while fossil fuel resources are depleting and GHG (greenhouse gas) emissions are threatening life on the Earth through global warming. At the United Nations Climate Change Conference (2010) in Cancun, Mexico, the international community recognized the need to keep warming below 2
C or the GHG concentration below 450 ppm CO2 equivalent (CO2-eq) in order to prevent the most catastrophic outcomes of global warming [1]. Based on this agreement, the primary goal of the IEA (International Energy Agency) is the reduction of energy-related CO2 emissions to below 50% of the levels in 2005 with renewables providing 40% of the primary energy supply by 2050 [2]. Renewable ethanol is a prominent solution to curb GHG emissions. If bioethanol were to replace petrol, GHG emissions would be reduced by more than 85% when considering the whole fuel cycle [3]. Global fuel ethanol production has increased from 13,096

* Corresponding author. Tel.: þ82 51 629 6453; fax: þ82 51 629 6429. E-mail address: [email protected] (J.J. Liu). http://dx.doi.org/10.1016/j.energy.2015.10.123 0360-5442/© 2015 Elsevier Ltd. All rights reserved.

million gallons in 2007 to 28,353 million gallons in 2011 [4] and will continue to increase in coming decades in order to help meet the IEA goals. However, further bioethanol production is limited by the availability of land and irrigation water. In addition, direct and indirect land use changes caused by crop cultivation and deforestation can result in a substantial carbon debt and high water consumption [5,6]. Seaweed, or macroalgae, as third-generation biomass, can circumvent these limitations. Because it grows in the ocean, seaweed does not depend on land availability, fertilizers, and irrigation for growth support, which eliminates the concerns regarding the scarcity of fertile land and its requirement for food supply. Another advantage of macroalgae (especially brown algae) over terrestrial biomass sources is the lack of lignin-type materials, which are resistant to biochemical conversion during biofuel production [7]. Fasahati et al. [8] evaluated the economy of bioethanol production from brown algae through a sugar platform. Their results showed that brown algae are a promising biomass for bioethanol production. However, the extent to which biofuels can replace petroleum-based fuels depends on the efficiency of the facility producing biofuels.

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Separation and dehydration of bioethanol from the fermentation broth is well-known to be an energy-intensive process that is usually responsible for a large portion of the energy requirements for the whole process [9,10]. Break-through separation technologies that reduce the capital costs, energy requirements, and carbon footprint are required. Membrane technology, which is less energy intensive and more environmentally friendly, is expected to play a crucial role in the transition to a sustainable chemical industry [11]. Pervaporation and vapor permeation are membrane separation processes with a wide range of uses in azeotropic systems, for solvent dehydration, and separation of organic mixtures [12,13]. Hybrid processes that combine pervaporation or vapor permeation with a traditional separation technique or with a chemical reactor are becoming increasingly common in industry, where traditional techniques perform poorly or are expensive [12]. Studies by Lipnizki et al. [14] and Sommer et al. [15] showed that combining distillation with pervaporation could lead to significant savings in investment and operating costs. Alzate and Toro [16] compared azeotropic distillation using benzene with a pervaporation system using multiple membrane modules for ethanol dehydration and concluded that the energy costs for pervaporation are approximately a fifth of those for azeotropic distillation. However, these authors considered cooling water instead of refrigerant for condensing permeates; water is a much cheaper coolant but is inapplicable for the temperature range within which condensation occurs on the permeate side. This is because, at vacuum pressures, the condensation temperature falls below 20
C depending on the degree of vacuum, and thus the use of a refrigerant is required for condensation [17]. Luis et al. [18] compared a hybrid distillation/pervaporation process with pressure swing distillation for separation of methanol and tetrahydrofuran. Their results showed that the hybrid process is superior from an environmental point of view. However, they did not provide an economic comparison of the methods, and, additionally, their simulations lacked heat integration, which directly € et al. [19] affects the energy requirements of the process. Niemisto performed a pilot-scale study of charcoal filtration and pervaporation for bioethanol dehydration using polyvinyl alcohol membranes. They found that filtration by activated carbon is an adequate pretreatment method for removing sulfur and other impurities, thereby enhancing the durability of pervaporation membranes. However, they did not evaluate the impact and costs of charcoal regeneration on the total energy requirements and economics of the process. In this study, we evaluate the sustainability of bioethanol production from brown algae, i.e., Laminaria japonica, through a VFA (volatile fatty acid) platform. In this platform, VFAs, including acetic acid, propionic acid, and butyric acid, are produced by anaerobic digestion of biomass using a mixed culture bacterial ecosystem. Subsequently, VFAs are recovered and hydrogenated to produce mixed alcohols [20e23]. The products of hydrogenation contain water, ethanol, propanol, and butanol. Considering the homogeneous minimum boiling azeotrope between water, ethanol, and propanol and the heterogeneous minimum boiling azeotrope between water and butanol, separation would be complicated and costly using distillation columns [24e26]. In this study, we propose two alternative processes for dehydration and recovery of alcohols: Hybrid pervaporation/distillation and hybrid vapor permeation/ distillation. The alternatives are evaluated based on their capital costs, MESP (minimum ethanol selling price), energy requirements, and carbon footprint. In addition, a sensitivity analysis is presented to identify the bottlenecks of each process, elaborate the main parameters affecting their economy, and reveal the uncertainties hidden in the processes.

2. Material and methods 2.1. Biomass, brown algae The brown algae family includes the largest and fastest growing seaweeds [27]. A total of 15.8 million wet tons of brown algae were harvested from wild habitats and aquaculture farms in 2010 [28]. Based on the data from the FAO (Food and Agriculture Organization) of the United Nations [28], the amount of masscultivated macroalgae (15.8 million wet tons) is four and six orders of magnitude greater than those of microalgae and lignocellulosic biomass (corn stover and switchgrass), respectively [29]. At present, the most promising brown algae species for biorefinery feedstock are L. japonica and U. pinnatifida, with annual productions of 5.1 and 1.5 million wet tons, respectively [28]. These two together account for over 40% of the total brown algae cultivated. Brown algae possess several advantages over terrestrial biomass that make it an interesting biomass for biofuel production [30,31]. Macroalgae can convert solar energy into chemical energy with a higher photosynthetic efficiency (6e8%) than terrestrial biomass (1.8e2.2%) [32]. They are capable of absorbing nutrients over their entire surface area [33] and can save energy due to zero requirements for internal nutrient transport [34]. In this study, we consider 400,000 ton/year dry brown algae L. japonica as feed for the plant. Brown algae are considered to contain 20% moisture upon arrival at the plant. The main carbohydrates of L. japonica include mannitol (a sugar alcohol), laminaran (a b-1,3-linked glucan that also contains mannitol), alginic acid (composed of mannuronic and guluronic acids), and fucoidan (a sulfated fucan that contains other sugars such as galactose, xylose, and uronic acid). Fasahati et al. [8] provided greater detail on the carbohydrate content of brown algae and their utilization by microorganisms for bioethanol production. McHugh [7] provided a comprehensive discussion of the available technology and methods for cultivation, harvesting, and processing seaweed. Table 1 shows the chemical composition and heating value of L. japonica brown algae. 2.2. Pervaporation and vapor permeation PV (pervaporation) and VP (vapor permeation) are membranebased processes in which a liquid (in the case of PV) or vapor (in the case of VP) is brought into contact with one side of a nonporous or molecularly porous membrane [36]. The driving force for both processes is the gradient of the chemical potential between the feed and permeate side, which can be intensified by increasing the feed temperature and flow-rate and lowering the pressure on the permeate side using vacuum pumps [10,37]. Most membranes are designed as hydrophilic or water permselective by incorporating attractive interactions between water and the membrane material, such as dipoleedipole interactions, hydrogen bonding, and ionedipole interactions [38]; this allows water molecules (kinetic diameter of 2.96 Å) to be selectively removed from larger molecules like ethanol (kinetic diameter of 4.5 Å), propanol (kinetic diameter of 4.7 Å), and 1-butanol (kinetic diameter of 5.0 Å) [39,40]. There are three categories of membranes: Polymeric, inorganic, and composite membranes [10]. Table 2 shows a literature survey of recent developments in PV and VP membranes for dehydration of ethanol. Polymeric membranes are based on organic polymer chains that are cross-linked to form tiny pores through which molecules can diffuse. The main polymeric membranes include poly(vinyl alcohol), chitosan, alginate, polysulfone, polyimides, polyamides, polyelectrolyte, and polyaniline membranes [12,41]. Polymeric membranes offer the advantage of lower cost. Nevertheless,

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Table 1 Chemical composition of the brown algae Laminaria species [35]. Proximate analysis

Dry basis, % w/w

Elemental analysis

Dry basis, % w/w

Ash Volatile solids (VS) Protein Lipids Mannitol Laminarin Alginates Cellulose Fucoidin Component Water Total solids

26 74 12 2 12 14 23 6 5 % wet basis 88 12

C H O N S P K

34.6 4.7 31.2 2.4 1 0.35 0.0096

Heating value HHV LHV

MJ/kg, dry basis 13.2 12.1

polymeric membranes are operated at somewhat lower temperatures, i.e., below that of the glass transition (<100
C), because of their insufficient thermal, mechanical, and chemical resistance [42]. However, the glass transition temperature and

characteristics of polymeric membranes, i.e., their selectivity, flux, and swelling, can be altered by blending and changing the degree of cross-linking and ratio of hydrophilic to hydrophobic moieties [12,43,44].

Table 2 Performance of pervaporation and vapor permeation membranes used in ethanolewater separation. %water Type Flux mkg2 h Selectivitya(a) T (
C) %Ethanol Polymeric membranes Polyimide 1 900 60 5/95 PAA/polyion complex 1.63 3500 60 5/95 PVA composite membrane 0.14 170 60 10/90 PERVAP®2201 0.1 100 60 10/90 PVAMMM membrane 0.5 1190 80 20/80 Sulfonated polysulfone membrane 0.87 500 35 10/90 Sodium alginate (Ca2þ)-polymer 0.23 330 50 10/90 ® Matrimid hollow fiber 0.16 130 45 15/85 Coated chitosan/cellulose acetate hollow fiber 0.23 23 25 10/90 Coated poly(vinyl alcohol)/poly sulfone hollow fiber 0.03 185 50 5/95 Grafted poly(acrylic acid)/polypropylene hollow fiber 0.20 11 24 30/70 Polyimide/Ultem® hollow fiber 0.49 124 60 15/85 Torlons e 4000T/Ultem® hollow fiber 0.66 50 60 15/85 Cellulose triacetate/Ultem® hollow fiber 1.28 466 50 15/85 ® PI/SPI/Ultem (3 wt%SPI) 3.20 55 60 15/85 PI/SPI/Ultems (4.5 wt%SPI) 3.80 21 60 15/85 PI/SPI/Ultems (3 wt%SPI) e thermal treatment 2.60 130 60 15/85 PI/SPI/Ultems (3 wt%SPI) e PDMS coating 2.70 104 60 15/85 PI/SPI/Ultem® (3 wt%SPI) e POSS modification 2.00 237 60 15/85 Inorganic membranes ECN silica 2.33 60 70 10/90 Mitsui zeoliteA 1.12 18000 70 10/90 Zeolite X 0.89 360 75 10/90 CHA-type zeolite membrane increased stability 2.89 >100,000 40 28/72 CHA-type zeolite membrane 4.14 39500 75 10/90 Ceramic membrane 0.458 724 87 46/54 Ceramic membrane 0.1 1633 79 5/95 NaA zeolite membranes a-Alumina (M-type) 2.2 10,000 75 10/90 Zeolite NaA 0.57 >10000 75 10/90 Mullite (M-type) (pervaporation) 2.1 42,000 75 10/90 Mullite (M-type) (Vapor permeation) 11 120 TiO2/steel (AS-type) 0.86 54,000 45 5/95 a-Alumina (AS-type) 0.25 8000 45 5/95 a-Alumina (M-type) 12.5 >5000 100 10/90 a-Alumina (M-type) (Vapor permeation) 10.5 >5000 125 10/90 a-Alumina (M-type) (Pervaporation) 5.6 10,000 75 10/90 a-Alumina (M-type) (Vapor permeation) 31 10,000 145 10/90 a-Alumina (AS-type) (Vapor permeation) 37 3900 145 10/90 a-Alumina (AS-type) (Vapor permeation) 20 4400 100 10/90 a-Alumina (AS-type) (Pervaporation) 8.5 10,000 75 10/90 NaA zeolite (Pervaporation) 5.9 9000 75 10/90 Zeolite NaA 2.1 2140 60 70/30 NaA 2.5 23000 75 10/90 NaA 3.80 3603 125 10/90 Mixed matrix membranes Hybsi® 1.70 139 70 5/95 PVA-KA zeolite mixed matrix 0.38 996 80 20/80 PVA-4A zeolite mixed matrix 1.50 530 80 20/80
 =YEth a Selectivity is defined as aw=Eth ¼ XYW , where X and Y are weight fractions of water and ethanol in the feed and permeate, respectively. W =X Eth

Ref [46] [47] [48] [49] [43] [50] [51] [52] [53] [54] [55] [56] [57] [58] [59] [59] [59] [59] [59] [45] [45] [60] [61] [62] [63] [63] [64] [64] [65] [65] [66] [67] [68] [68] [69] [69] [70] [70] [70] [71] [72] [73] [74] [75] [43] [76]

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Inorganic membranes fabricated from ceramics or zeolites have superior temperature stability, solvent resistance, and mechanical strength than polymeric membranes [45]. They also provide higher permeate flux and selectivity than polymeric membranes (Table 2). Ceramics are hard materials that offer good mechanical stability with high melting points of over 1000
C, and they are able to operate over a wide pH range in any organic solvent [12]. Zeolites (aluminosilicates) offer a good basis for a separation material because of their highly ordered, well-defined structures. Depending on the type of zeolite and ratio of aluminum, oxygen, and silicon in its framework, a wide number of zeolite structures can be synthesized with pore sizes ranging from about 3 to 8 Å [12]. However, inorganic membranes are often expensive and more difficult to produce on an industrial scale [12]. Commercialized zeolite membranes suffer from dealumination of the zeolite framework in acidic solutions, with consequent degradation of the zeolite structures. Inorganic silica is known to be hydrothermally unstable, which becomes apparent at temperatures as low as ~70
C [77]. Upon exposure to moisture, hydrolysis and net transport of silicon species lead to substantial loss of permeability within hours. After prolonged exposure to water, dense silica particles with associated large pores are formed, resulting in a loss of selectivity [78,79]. Composite membranes, which are also known as mixed matrix membranes, have been developed to overcome the drawbacks associated with inorganic and polymeric membranes. Their structure consists of a polymeric base membrane in which an inorganic material is dispersed and locked into the polymer matrix. The addition of the inorganic material can help strengthen the mechanical properties of the membrane and reduce the free volume through which molecules may diffuse.

2.3. Process description Fig. 1 shows the general configuration of a plant for mixed alcohol production from brown algae. The process units include anaerobic digestion (A), VFA recovery (B), hydrogenation (C), and

mixed alcohols recovery (D, E, and F). As shown in Fig. 1, three alternatives are considered for mixed alcohols recovery: Classical molecular sieves and distillation (D), hybrid PV and distillation (E), and hybrid VP and distillation (F).

2.3.1. Anaerobic digestion In the AD (anaerobic digestion) unit, brown algae are fermented to produce volatile fatty acids. The volatile fatty acids include acetic acid, propionic acid, and butyric acid. Pham et al. [80] performed an experimental study on the anaerobic digestion of L. japonica and reported a VFA yield of ~0.307e0.412 g VFA/g dry L. japonica after five days of fermentation at 35
C. Therefore, in this study, an average yield of 0.35 g VFA/g dry brown algae and a five-day fermentation time are considered. Their results showed a molar product distribution of 67% acetic acid, 22% propionic acid, and 11% butyric acid. Fasahati and Liu reported the fermentation reactions and conversions required to reach this product selectivity and yield [81]. Tables Se1 of the supplementary data shows the stoichiometric reactions and conversions that are assumed to occur in fermentation reactors. Fresh water, along with recycled water from the MA (mixed alcohols) recovery unit, is added to the AD reactors to obtain a 20% solid loading. LP (low pressure) steam is added to warm the feed to an operating temperature of 35
C. Iodoform or bromoform can be added as an inhibitor to the methane production pathway. Granada et al. [82] reported that an inhibitor concentration of 30 ppm prevents methane production; therefore, this value is adopted in this study. In the AD reactors, chilled water is considered as the coolant to maintain the AD temperature at 35
C. A VFA concentration in the product stream of 5% is considered [80,82]. The products of AD are separated in a solid/liquid separator. The solid residue (digestate) of fermentation contains a considerable amount of nutritive components, such as nitrate and phosphate, fibers, and some carbon compounds. Therefore, the digestate can be a useful fertilizer. About 20% of the digestate is recycled back to the fermentation reactors as an inoculum and a nutrient source [83]. The remaining

Fig. 1. General configuration for mixed alcohol production from brown algae: Anaerobic digestion (A), VFA recovery (B), hydrogenation (C), and mixed alcohol recovery using molecular sieves and distillation (D), hybrid PV and distillation (E), and hybrid VP and distillation (F).

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assessed by performing a sensitivity analysis in Section 3.4.3. The aqueous phase (raffinate) in the extraction tower is saturated with MTBE. Consequently, MTBE is recovered in a downstream stripping tower. LP (low-pressure) steam at 164
C is used to strip MTBE from the water. The flow-rate of steam is adjusted to recover 99% of the MTBE from the top of the stripping column. The water at the bottom of the striping column is considered to be waste water. The extract stream from the extraction column is sent to a rectification column for separation of VFA from MTBE and a small amount of dissolved water. The extraction agent and water accumulate at the top end of the rectification tower and the VFAs accumulate at the bottom end, resulting in a VFA concentration of ~99.99%. MTBE recovered from the top of the stripping and rectification columns is condensed and sent to a decanter. The organic phase containing mostly MTBE is recycled to the extraction column, and the water phase is sent to the stripping column.

digestate is considered as a fertilizer that can be sold on the local market. The impact of the digestate selling price on the MESP is studied in the sensitivity analysis. The separated liquid stream, which contains mostly water and VFAs, is sent to the VFA recovery unit. 2.3.2. VFA recovery In the VFA recovery unit, VFAs are dehydrated and recovered via extraction and distillation. According to reference [84], VFAs can be easily recovered at a purity above 99.99 wt% because of the wide boiling temperature range between acetic acid and MTBE (methyl tert-butyl ether) and their large separation factor. Fasahati and Liu [81] evaluated the application of a hybrid PV and extraction/ distillation process for the recovery of VFAs. However, their results showed that, at low VFA concentrations, the PV process requires a huge membrane area for separating water and refrigeration energy for condensation of permeates. Therefore, the process becomes very expensive in terms of energy and capital costs. The state-ofthe-art technology for VFA recovery is azeotropic rectification, with or without an extraction stage depending on the VFA concentration. In the case of acetic acid concentrations below 40 wt%, the acetic acid is initially extracted from the aqueous solution with a suitable extraction agent before pure recovery occurs during rectification of the azeotropic mixture [84]. Table 3 shows several potential extraction agents that are suitable for the recovery of VFAs. During selection of the extractive agents, several characteristics must be taken into account, such as the solubility in water, absorption capacity, distribution coefficient, price, availability, and composition of the azeotrope, and the requirements in terms of environmental and health protection. The lower vaporization enthalpy of azeotropic mixtures results in lower energy consumption during the recovery process. Thus, the use of EtAc (ethyl acetate) or MTBE (methyl tert-butyl ether) as extraction agents can reduce the energy consumption. In this study, MTBE is selected as an extractive agent because its lower boiling point (55
C) and enthalpy of vaporization (322 kJ/kg) are advantageous over those of EtAc [84]. The density of the VFA mixture is higher than that of MTBE. Therefore, VFA is fed into the top end of the extraction tower and MTBE is added from the bottom. Inside the tower, the MTBE flows towards the top of the column and selectively extracts VFAs from the water. The extraction ability of MTBE for VFAs increases in the following order: Butyric acid > propionic acid > acetic acid [85]. By increasing the flow-rate of MTBE relative to the flow-rate of VFAs, a larger amount of acetic acid can be extracted from the fermentation products. However, higher acetic acid recovery increases the energy and capital costs for handling and separation of MTBE from recovered VFAs. The results of Aspen Plus simulations showed that acetic acid recovery above 95% strongly increases the MTBE flow-rate and energy costs. Therefore, in the base case, the flow-rate of MTBE is adjusted to recover 95% of the incoming acetic acid in the extract phase. This eventually results in almost 100% recovery of propionic and butyric acid. The impact of acetic acid recovery on the process economics is

2.3.3. Hydrogenation Carboxylic acids can be hydrogenated to produce alcohols. The process demands activation of the carbonyl bond in the carboxylic acid followed by the addition of hydrogen. The main challenge of the process is activation of the VFA carbonyl group, which is very stable because of its weak polarizability. Thus, expensive hydride agents, such as LiAlH4, are normally required to reduce carboxylic acids to alcohols. These reactions are not economical for industrial use [86,87]. Several alternative processes have been developed for hydrogenation of VFAs including catalytic conversion in the vapor phase [88,89], liquid phase conversion [90], hydrogenation of esterified carboxylic acids [91,92], and biochemical conversion [93,94]. The current industrially practiced technology for VFA hydrogenation is a patented process developed by Schuster et al. [95]. This is a gas phase process with a high alcohol yield of 97% [95]. Gasphase hydrogenation requires high pressure and temperature, and a predominantly cobalt-containing catalyst is used. The reaction temperature and pressure are 290
C and 60 bar, respectively. Based on the study by Schuster et al. [95], liquid acetic acid is extremely aggressive and corrosive and damages the catalyst, while gas-phase acetic acid is not corrosive. Consequently, the gas-phase reaction not only leaves the catalyst undamaged, but also makes it unnecessary to use expensive metals or metal alloys as construction materials for the reactor and its accessories [20]. Eq. (1) shows the hydrogenation reaction for acetic acid as a representative component:

C2 H4 O2 þ 2H2 /C2 H6 O þ H2 O:

(1)

Propionic acid and butyric acid follow a similar stoichiometric hydrogenation reaction. Hydrogen (based on a H2:VFA molar ratio of 2.1:1) is compressed in a three-stage compressor, with inter-stage cooling to 60
C by cooling water, and sent to the reactor [20]. A catalyst consumption rate of 460 mg catalyst/kg mixed alcohols during

Table 3 Extraction agents for the separation of acetic acid from water [84]. Name

Ethyl acetate C4H8O2 (EtAc) Isopropyl acetate, C5H10O2 (iPrAc) n-Propyl acetate, C5H10O2 (nPrAc) Methyl propyl ketone, C5H10O (MPK) Methyl isobutyl ketone, C6H12O (MIBK) Methyl tert. butyl ether, C5H12O (MTBE)

Average distribution coefficientkg/kg

0.84 0.55 0.50 0.97 0.50 0.75

Densitykg/m3

900 877 891 810 810 740

Enthalpy of vaporizationkJ/kg

395 361 336 384 488 322

Boiling point
C

76.7 88.6 101.6 102.3 115.9 55

Azeotrope WaterWt %

T
C

8.47 10.50 13.20 19.50 24.30 4

70.4 76.5 82.2 83.3 87.9 52.6

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hydrogenation is considered [20]. The VFA stream is heated to the hydrogenation reaction temperature of 290
C using a combination of preheating by heat exchange with the products of the hydrogenation reactor and a fired heater. The heat from the hydrogenation reaction and the hot effluent of the hydrogenation reactor are used to produce LP steam from the BFW (boiler feed water) and preheat the inlet stream. The products are further cooled to 65.5
C by means of cooling water before being sent to the first flash tank. From the first flash tank, 98% of the recovered vapor phase, containing mostly hydrogen, is compressed in a multi-stage compressor and recycled to the hydrogenation reactor, while the remaining 2% is purged. The bottom stream of the first flash drum is further flashed to 2.4 bar and sent to the second flash drum [20]. The vapor from the second flash drum and purge gases from the first flash drum contain primarily unreacted hydrogen and ethanol and are used as fuel in the fired heater before the hydrogenation reactor. The bottom stream of the second flash drum, which contains a mass fraction of 25% water, 42% ethanol, 19% propanol, and 12% butanol, is sent to the alcohol recovery unit for separation of water, ethanol, and heavier alcohols. 2.3.4. Mixed alcohol recovery The mixed alcohol recovery unit first dehydrates the wet alcohols, and then separates the dry mixed alcohols to fuel grade ethanol (99.5%) and heavier alcohol co-products including propanol and butanol. Water forms a homogeneous minimum boiling azeotrope with ethanol and propanol and a heterogeneous minimum boiling azeotrope with butanol; this makes separation of the mixture very complex and difficult to achieve using distillation. Therefore, water is first removed from the mixed alcohols by molecular sieves, pervaporation, or vapor permeation, and then the dried mixed alcohols are sent to the distillation column to separate ethanol from propanol and butanol [24e26]. 2.3.4.1. Molecular sieves and distillation (classical). Molecular sieves and distillation is a classical process that its general structure has been considered in NREL (National Renewable Energy Laboratory) studies [24,25], which use molecular sieves to dehydrate wet alcohols and an alcohol distillation column to separate dry alcohols. In NREL's studies, the products of regeneration of the molecular sieves are recycled, mixed with syngas, and transferred to an alcohol synthesis reactor. However, in this study, we added another column, i.e., an azeotropic column, to separate the ethanol and water. MA (mixed alcohols) from the hydrogenation unit are preheated by the bottom products of the alcohol distillation column and further heated using LP steam for evaporation and superheating to 116
C before being fed to the molecular sieves [96]. Water is removed using vapor-phase molecular-sieve adsorption, and the dried mixed alcohols are sent to the alcohol distillation column where ethanol is separated from propanol and butanol. The ASTM ethanol fuel specifications require a purity of 99.5% for ethanol. Therefore, 99% of the incoming ethanol is recovered from the top of the column with a mass purity of 99.5%. Propanol and butanol are recovered from the bottom of the column and are considered to be co-products of the plant. It is anticipated that the heavier alcohols may find use as excellent gasoline additives or gasoline replacements [26]. About 65% of the vapor product ethanol is used to back-flush the molecular sieves by applying a vacuum during its regeneration [96], and the rest is liquefied as the ethanol product of the plant. The product of molecular sieve regeneration contains a mixture of ethanol and water, which is liquefied using cooling water and sent to an azeotropic column. Prior to entering the column, the mixture is preheated to 65
by heat from the product ethanol, bottom

product of the alcohol separation column (heavier alcohols), and bottoms of the azeotropic column. Due to the ethanol-water azeotrope, pure water can be recovered from the bottom of the column and is recycled to the fermentation unit. The top product of the azeotropic column contains water and ethanol at the azeotropic point (92.5 wt% ethanol) and is recycled to the molecular sieves. 2.3.4.2. Hybrid PV/distillation (PV). In the hybrid PV/distillation process, water is removed using PV membranes and the dehydrated alcohols are sent to distillation columns for separation of ethanol from propanol and butanol. The membrane performance reported by Sato et al. [70] for an a-alumina NaA zeolite membrane given in Table 2 is considered in this study. The membrane enables a permeate flux of 8.5 kg/m2 h at 75
C with a selectivity greater than 10,000. The wet alcohols received from the hydrogenation unit are preheated, first using the bottoms of the alcohol distillation column and then by LP steam, to the PV temperature of 75
C. The energy to evaporate the water in PV membranes is obtained from cooling the liquid stream. Thus, reheating between modules is required to keep the temperature constant at 75
C; this is achieved using steam-heated heat exchangers. The permeate gases are condensed using liquid propane at 20
C. The vacuum pumps are considered to provide a constant pressure of 1 kPa on the permeate side. The liquefied permeate stream contains pure water and is used to cool the ethanol product of the alcohol distillation column before recycling to the pretreatment unit. The retentate of the PV membranes containing dry alcohols is sent to the alcohol distillation column to recover 99% of the ethanol at a purity of 99.5% from the heavier alcohols that include propanol and butanol. The top and bottom products of the column are cooled by heat integration and cooling water to 38
C as products of the plant. 2.3.4.3. Hybrid VP/distillation (VP). Hybrid VP/distillation has a similar structure as the hybrid PV/distillation process except that the heat integration structure of the VP process is different. The membrane considered for VP in this study is the a-alumina NaA zeolite membrane developed by Sato et al. [70]. The membrane allows a permeate flux of 37 kg/m2$h and a selectivity of 3900 (Table 2). The wet alcohols are preheated and evaporated via heat exchange with the bottoms from the alcohol distillation column and retentate vapors and with MP (medium-pressure) steam to reach a temperature of 145
C before contacting the VP membranes. Vacuum pumps are used to maintain a constant pressure of 1 kPa on the permeate side. Liquid propane at 20
C is used to condense the permeate vapors. The liquefied permeate stream contains pure water and is used to cool the ethanol product stream before recycling to the pretreatment unit. The bottoms of the alcohol distillation column contain heavier alcohols that are cooled to 38
C by heat exchange with the wet alcohols and then with cooling water. 2.4. Process simulation Aspen Plus v.8.4 (Aspen Tech, Cambridge MA) is used to simulate the process. Due to the existence of VFAs and the non-ideal behavior of the liquid phase, the non-random two-liquid NRTLHOC thermodynamic/activity model is selected as the property package. The method uses the Hayden-O'Connell equation of state as a vapor phase model. The chemical composition of brown algae reported by Reith et al. [35] presented in Table 1 is used as the biomass composition. Most of the components are available in the Aspen Plus data banks. Fasahati et al. [8] defined the component properties for the unavailable components using the physicalproperty database developed by the NREL [96,97]. Pumps are modeled with an assumed efficiency of 80% in the simulation. The hydrogenation compressor is modeled as a three-stage centrifugal

P. Fasahati, J.J. Liu / Energy 93 (2015) 2321e2336

compressor with 78% polytropic efficiency and inter-stage cooling to 60
C [24]. The molecular sieves and membranes are modeled as component separators in Aspen Plus. A calculator block is used to calculate the composition of the permeate based on the specified selectivity and VFA concentration in the retentate. The RadFrac module of Aspen Plus is employed for rigorous mathematical analysis of all distillation columns. Design specifications within RadFrac are used to achieve the objective purifications and recoveries in each column. In terms of the molar split fractions, optimization of utility consumption, and thermal integration of the columns, the design and modeling of each column is performed according to Smith et al. [98]. Guidelines regarding the number of stages, type of trays, reflux ratio, column pressure, etc., are extracted from references 24, 96, and 99. 2.5. Heat integration Heat integration is performed separately for each process in order to minimize the utility requirements. An inventory of the heat demand and supply within the plant is made and ordered by temperature range. Hot and cold streams are matched under the following criteria: Firstly, matches between near streams (same plant section) are preferred; secondly, respective minimum temperature differences of 20
C, 30
C, and 40
C are imposed for liquideliquid, liquidegas, and gasegas heat transfer matches [99]. The bottom fractions of the distillation columns are used to preheat the feed streams of the column. The heat produced during the hydrogenation reactions is recovered by producing an LP steam from the BFW. The cold permeate water (~7
C) of the membranes is used as a coolant before recycling to the fermentation unit. Heating and cooling demands are covered by process-to-process heat transfer unless greater energy demands are needed, in which case appropriate utilities are applied. LP steam at 164
C and MP steam at 186
C are used as heating sources. Cooling demands above 40
C are met by cooling water. Cooling below 40
C is only required in the AD fermenters and for liquefaction of the permeate, where chilled water (4
Ce15
C) and liquid propane (20
C) are used, respectively. 2.6. CO2 emissions Since the three process alternatives have different utility demands, their CO2 emissions during operation are different. PV and VP require cooling water in the condenser of the alcohol separation column and refrigerant for condensing the permeates, while the classical process requires only cooling water for its cooling demands. In addition, PV and VP demand more electricity than the classical process because of the power requirements of vacuum pumps. VP operates at a higher temperature of 145
C compared to PV, which operates at 75
C; therefore, VP requires MP steam for provision of heat while LP steam is sufficient to satisfy the heat demands of the PV and classical processes. It is of interest to identify the process that has the lowest CO2 emissions during its operation. It is noteworthy that the analysis of the CO2-eq emissions is solely related to direct emissions during operation, and a complete life-cycle analysis can provide more complete results regarding the relative merits of the technologies. The approach used to calculate CO2 emissions is based on a methodology defined by the US EPA (Environmental Protection Agency). The following assumptions are made as a basis for the calculations:  The steam-production GHG emissions are calculated based on the method given by the EPA considering natural gas as the fuel source and a boiler efficiency of 80% [100]. Subsequently, the





 

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GHG emissions from burning natural gas are extracted from Ref. [101]. CO2 emissions due to electricity consumption are calculated based on average US annual output emission rates for electricity production [102]. CO2 emissions caused by refrigerants are derived by calculating the amount of electricity required to produce one ton of refrigeration. Based on Turton et al. [103], compression refrigeration with a 38
C condenser requires 1.31 kW/ton of refrigeration at 20
C and a ton of refrigeration is the removal of 12.7 MJ/h of heat. The electricity requirements for the production of 1 m3/h cooling water is 0.254 kW according to reference [96]. Emission factors are converted to CO2-eq using 100-year GWPs (global warming potentials). The 100-year GWPs for CH4 and N2O are 25 and 298, respectively [101].

2.7. Techno-economic model The techno-economic model considered in this study is similar to that developed by the NREL for bioethanol production from lignocellulosic biomass [96]. The TCI (total capital investment) is first computed from the total equipment cost. Next, the variable and fixed operating costs for each process are determined. Based on these costs, a discounted cash flow analysis is developed to determine the MESP required to reach a break-even point after 20 years of plant life considering a 10% internal rate of return and 100% equity financing. The techno-economic model is based on nth-plant economic assumptions; these include that the economic analysis does not describe a pioneer plant but rather that several plants using the same technology have already been built and are operating. This is a relatively realistic assumption since the process units have been separately validated at lab and industrial scales. Therefore, the process is believed to have a high TRL (technology readiness level) based on the NASA (National Aeronautics and Space Administration) definition [104]. Terrabon Inc. has constructed an AD unit for VFA production (50,000 gal/year) from biomass [105]. De Dietrich Process Systems GmbH has designed and industrially practiced the recovery of acetic acid from aqueous solutions [84]. Hydrogenation of VFAs to alcohols [20,95] and separation of water alcohol mixtures are also well-developed industrial processes [106]. As a result, the process units are mature technologies that allow implementation of the process of mixed-alcohols production from brown algae at an industrial scale and reach a TRL9 within short period. Table 4 shows the economic parameters for the discounted cash flow analysis. In NREL's economic model, a three-year construction period, 8410 operating hours per year, and 30 years of plant life are assumed. However, in this study, the construction period, operating hours, and plant life are assumed as follows based on the economic model developed by Perales et al. [99]: One-year construction period, 8000 operating hours per year, and 20 years of plant life. The longer construction period for NREL's process model is mainly because of their larger plant scale (700,800 ton/year) and additional waste water treatment unit, combustor, boiler/turbogenerator, utility units, etc., in their bio-refinery structure. The impact of the longer plant life is evaluated for all the processes in the sensitivity analyses in Section 3.4.1. Based on Perry's Chemical Engineers' Handbook [107], the startup time for a moderately complex plant should be about 25% of the construction time. Therefore, a 3-month start-up period is assumed for the process, which is similar to NREL's assumption. The MACRS (modified accelerated cost recovery system) reported by the U.S. IRS (Internal Revenue System) is used to

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Table 4 Techno-economic model parameters for discounted cash flow analysis. Parameter

Value

Cost basis year Plant life Internal rate of return (IRR). Depreciation method (recovery period) Tax rate Working capitala Landa Salvage value Construction period Startup period Revenues during start-up Variable costs incurred during start-up Fixed costs incurred during start-up Operating hours per period

2012 dollars 20 years 10% 200% declining balance (7 years) 35% per year 5% of fixed capital investment (FCI) 6% of installed cost 0 M$ One year 3 Month 50% 75% 100% 8000 h/year

a

Working capital and cost of land are recovered at the end of plant life.

determine the federal taxes to be paid. The model considers zero salvage value at the end of the plant life. According to IRS publication 946, algae-based bioethanol plants fall under asset class 49.5, i.e., “waste reduction and resource recovery plants” [108], which uses a seven-year recovery period. The process equipment are sized based on the mass and energy balance results of the simulation. The capital costs of the azeotropic column, molecular sieves, and alcohol separation column are scaled based on the NREL's design reports using Eq. (2) [96,109]. Tables Se2 in the Supplementary Data shows a summary of the equipment cost, scaling exponent, year of quote, and installation factors obtained from NREL's reports.


CostNew ¼ Costbase

ScaleNew Scalebase

n (2)

The values of the scaling exponents (n) are reported for equipment in Tables Se3 of the Supplementary Data. These values are also used to evaluate the impact of the plant scale on process economics in the sensitivity analyses. The extraction column, rectification column, and decanter in the VFA recovery unit are selected and designed based on the criteria given in the literature [107e112]. Further details about equipment design can be found in the report by Fasahati and Liu [81]. After sizing the equipment, the cost estimation for the equipment is performed based on a modulecosting technique and relationships published by Turton et al. [103]. The base year for the economic analysis is 2012. Therefore, all the equipment costs are updated to 2012 dollars using the CEPCI (chemical engineering plant cost index) through the following equation:

2012 cost ¼ ðbase costÞ

CEPCI2012 CEPCIbase year

3. Results and discussion The three processes are simulated based on the methodology explained in the previous sections, and the total capital investment is calculated. The simulation results are shown in Table 6. The results show that the classical process requires 25% more cooling water and 23% more LP steam than the PV and VP processes. However, the PV and VP processes require 2.61 Gcal/h of refrigerant to condense the water permeates of the membranes. The ethanol production of the classical process is ~138 kg/h less than those of the PV and VP processes; this is due to the criterion considered in this study to recover 99% of ethanol from the alcohol separation column at a purity of 99.5%. Since, in the classical process, part of the ethanol product is used to regenerate the molecular sieves, the total flow-rate of ethanol circulated between the membrane and Table 5 Data for calculating operating costs. Ref.

! (3)

The price of vacuum pumps is generally estimated based on Eq. (4), as reported by Ryans and Croll [113] in the year 1981, as follows:

CRotary Blower ¼ 22000ðHP=10Þ0:4 ;

per m2 by the membrane area required. The total equipment costs for the membrane module, including the casing, sealing, instrumentation, etc., are calculated by applying a 10% overhead rate on the membrane costs. The FCI (fixed capital investment) is calculated by adding the direct and indirect costs. The TCI (total capital investment) is calculated as the sum of the working capital, land, and FCI. Tables Se4 of the Supplementary Data presents the cost factors for estimating the direct and indirect costs. The operating costs include the variable operating costs (e.g., processing chemicals) and fixed operating costs (e.g., employee salaries, overhead, maintenance, and insurance). The NREL's economic model calculates the labor costs by considering the recommended number of employees and associated salaries. In this study, the labor cost is assumed to be 1.6% of the total installed costs based on the economic model by Perales et al. [99]. The basis for the fixed and operating cost calculations are given in Table 5. The MESP is affected by the revenues from the co-products of the plant. These co-products include the digestate, which is the solid residue from the AD, and heavier alcohols produced in the plant, i.e., propanol and butanol. The demand for the digestate can be limited in the local market, thereby affecting the selling price. A digestate selling price of $10.4/ton was projected by Dave et al. [83]. The selling price of the heavier alcohols is assumed to be the market price of n-butanol, i.e., $1.15/kg [116]. Sensitivity analyses on the selling price of the co-products are performed to predict the impact on the plant economics and MESP.

(4)

where HP is the motor horsepower of the vacuum pump. The membrane costs are derived from the 2015 target of the National Energy Technology Laboratory of the U.S. Department of Energy (US-DOE-NETL) [114] for hydrogen production using ceramic and metallic membranes, which is roughly $1000/m2. The durability or membrane lifetime is another uncertainty. The 2015 target of the DOE is five years, which is adopted in this study [114,115]. The membrane costs are calculated by multiplying the membrane costs

Fixed operating costs Labor cost Maintenance cost Property insurance and tax Variable operating cost Biomass Cost of macroalgae cultivation Transport cost Total macroalgae cost Hydrogenation catalyst LP steam at 164
C MP steam at 185
C Waste water Hydrogen Cooling water Chilled water Refrigerant at 20
C MTBE Process water

1.6% of total installed cost 3.0% of total installed cost 0.7% of FCI

[99] [96] [96]

Laminaria japonica 54.4 $/ton (dry basis) 13.6 $/ton (dry basis) 68 $/ton (dry basis) 18.975 $/lb catalyst 12.68 $/ton 13.71 $/ton 0.041 $/m3 1.5 $/kg 0.013 $/ton 1 $/ton 7.89 $/GJ 1100 $/ton 0.27 $/ton

[83]

[20] [103] [103] [103] [114] [117] [117] [103] [118] [119]

P. Fasahati, J.J. Liu / Energy 93 (2015) 2321e2336 Table 6 Results of simulation for PV, VP, and classical processes.

Brown algaea Ethanol Mixed alcohols (C3 and C4) Digestate Fresh water BFW Waste water LP steam MP steam Refrigerant (Gcal/hr) Chilled water Cooling water Electricity (kW) Hydrogen MTBE Inhibitor a

PV (kg/hr)

VP (kg/hr)

Classical (kg/hr)

62,500 7415 6077 24,831 315,687 5461 383,455 131,613 e 2.61 629,470 5,862,120 932 1006 37 10.2

62,500 7414 6090 24,831 315,699 5462 383,454 126,336 10,295 2.61 629,469 5,878,670 930 1006 37 10.2

62,500 7276 6202 24,831 314,443 5458 381,862 161,884 e e 629,428 7,330,330 920 1006 32 10.2

Including 20% moisture.

alcohol separation columns is larger than that for the PV and VP processes. Therefore, the 1% of ethanol that is not recovered from the top is a larger quantity than those for the PV and VP processes. This consequently results in a higher flow-rate of heavier alcohols in the classical process than in the PV and VP processes.

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The classical process also has a higher total fixed operating cost and a variable operating cost that is higher by about $2 million/year. The higher utility demand is the main reason for the difference in variable operating costs. The small difference in total fixed operating costs between PV and VP is due to the higher maintenance costs for the larger membrane area required for the PV process.

3.2. Energy consumption The results of the simulations are used to obtain the energy requirements for each unit. Later, the energy costs are calculated based on the flow-rate and utility costs given in Table 5. Fig. 2 shows the utility cost breakdown for the fermentation, VFA recovery, and hydrogenation units. The fermentation and VFA recovery units have utility costs of $5.2 and $13.2 million/year, while the hydrogenation unit provides positive revenues of about $0.22 million/year by producing 5.46 ton/h LP steam from BFW using heat from the hydrogenation reactors and its hot products. The hydrogenation unit requires more electricity input than the other units because of the power requirements of the hydrogen compressors. As can be seen from Fig. 2, the main utility cost in the fermentation unit is chilled water ($5 million/year), which is required to keep the fermentation temperature constant at 35
C. The VFA recovery unit is the main consumer of LP steam. Considering the low VFA concentration (5%) and high volumetric flow-rate of the

3.1. Total capital investment Table 7 shows the total capital investment and operating costs for each process. The capital costs of the anaerobic digestion, VFA recovery, and hydrogenation units are identical for each process, but vary for the alcohol recovery unit. The VFA recovery unit has higher capital costs than the other units because it receives 348 ton/ h fermentation broth, and, after recovery and purification, delivers 17 ton/h VFA to the hydrogenation unit. The reduction in material flow-rate results in smaller equipment sizes and thus lower capital costs for the hydrogenation and alcohol recovery units. From Table 7, the capital cost of the alcohol recovery unit for the classical process ($8.0 million) is more than three times greater than that of the PV and VP processes, which have similar capital costs ($2.6 and $2.4 million, respectively). However, since the permeate flux in VP is larger than in PV, less membrane area is required for water removal in the VP process. This results in a negligible decrease in the capital cost of the alcohol dehydration unit for the VP process compared to PV. The higher total installed costs of the classical process consequently result in higher land, working capital, and TCI required than required for the PV and VP processes. The TCI for the classical process is about $11 million more than those of the PV and VP processes.

Fig. 2. Utility-cost breakdown of the fermentation, VFA recovery, and hydrogenation units.

Table 7 Plant cost worksheet for total capital investment (TCI), total variable and fixed operating cost (million$).

Fermentation VFA recovery Hydrogenation Alcohols recovery Total installed costs (TIC) Total direct costs (TDC) Total indirect costs Fixed capital investment (FCI) Land Working capital Total capital investment (TCI) Total variable operating cost (per year) Total fixed operating cost (per year)

PV (million $)

VP (million $)

Classical(million $)

9.4 16.2 4.1 2.6 32.3 37.9 22.8 60.7 1.9 3.0 65.7 67.7 1.98

9.4 16.2 4.1 2.4 32.0 37.6 22.6 60.2 1.9 3.0 65.2 68.3 1.90

9.4 16.2 4.1 8.0 37.7 44.3 26.6 70.9 2.3 3.5 76.7 70.2 2.22

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fermentation products, a large amount of extractive agent (i.e., MTBE) and steam is required to recover VFAs from water. The steam is required in the stripping column and reboiler of the VFA distillation column. In order to compare the energy costs of the PV, VP, and classical processes, the energy costs of the alcohol recovery unit are shown separately in Fig. 3. The utility costs of this unit for the PV, VP, and classical processes are $2.0, $2.6, and $4.6 million/year, respectively. The PV process has the lowest energy requirements. The VP process operates at a higher temperature of 145
C and therefore requires MP steam, which increases the total energy requirements. In addition, in the VP process, the feed stream must be completely evaporated before contacting the membrane, while in the PV process, heat provision is required only for evaporation of the permeate. Thus, the energy demands of the PV process are lower than those of VP. The classical process is similar to the VP process in that the wet mixed alcohols require evaporation before they are fed to the molecular sieves. Utilization of part of the ethanol product for regeneration of the molecular sieves also increases the energy demands and requires an additional distillation column for recovering ethanol from the azeotropic mixture with water. In this way, a mixture of ethanol and water is continuously circulated in the classical process, which results in larger equipment and higher energy demands. 3.3. MESP (minimum ethanol selling price) The MESPs of the PV, VP, and classical processes are calculated to be $1.06/gal, $1.08/gal, and $1.24/gal, respectively. Fig. 4 shows the MESP breakdown of the three process alternatives and contributions of different parameters to the total MESP. The PV and VP processes have similar MESP values, with the increased utility costs of VP balanced by its lower fixed operating and capital costs. As was explained in previous sections, the classical process has both higher utility and capital costs than the hybrid processes. Therefore, the MESP of the classical process is ~15% higher than those of the PV and VP processes. The MESP values obtained for the different processes in this study are comparable to those reported by Phillips et al. [24] for indirect gasification and Dutta and Phillips [25] for direct gasification of lignocellulosic biomass (forest residues); they reported

Fig. 3. Utility-cost breakdown of the alcohol separation and recovery units for hybrid PV/distillation, hybrid VP/distillation, and classical molecular sieve/distillation processes.

Fig. 4. Minimum ethanol selling price ($/gal) for hybrid PV/distillation, hybrid VP/ distillation, and classical molecular sieve/distillation processes.

respective MESPs of $1.29/gal and $1.95/gal for the indirect and direct gasification processes using their 2007 cost estimations and 2012 performance targets, respectively. The MESP calculated for the PV process in this study shows that ethanol production from brown algae through a VFA platform can be cheaper than thermochemical bioethanol production from second-generation biomass. 3.4. Sensitivity analysis 3.4.1. Impact of economic parameters Because of hidden uncertainties in the processes and economic parameters, it is of interest to perform a single-point sensitivity analysis to evaluate the impact of key parameters on the MESP. In a single-point sensitivity analysis, one parameter at a time is varied between ±25% and ±50% of its base value (except for the alcohol co-product selling price, which varies between ±25% of its base value), and the impact on the MESP is calculated using the techno-economic model. The values on the x-axis are normalized to allow a comparison between different units. The intersection of lines on the charts indicates the MESP default values used in the base cases (x-axis: 0%). Figs. 5e7 show the results of the sensitivity analyses for the PV, VP, and classical processes, respectively. The results of the sensitivity analysis show that, in the order of their magnitude, the mixed alcohol co-product value, biomass price, steam price, FCI, and finally IRR (internal rate of return) have the greatest impact on the MESP for all processes. The MESP changes linearly as the parameters are varied between their minimum and maximum values, except for plant life, for which the MESP shows higher sensitivities at lower plant life than at higher plant life. As expected, the mixed alcohol co-product valuation has the largest effect on the MESP, since a significant quantity of co-product is produced by the plant. Varying the selling price of the mixed alcohol

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Fig. 5. Impact of changing key parameters from their base values for the PV process.

co-product by ±25% of its base value of $1.15/gal results in changes of the MESP from $1.77/gal to $0.35/gal, $1.78/gal to $0.37/gal, and $1.97/gal to $0.51/gal for the PV, VP, and classical processes, respectively. The second largest impact is caused by variations in the biomass price given that varying the biomass price by ±25% from $51/ton to $85/ton varies the MESP from $0.64/gal to $1.49/gal, $0.65/gal to $1.50/gal, and $0.81/gal to $1.67/gal for the PV, VP, and classical processes, respectively. Similarly, a ±50% change in biomass price from $34/ton to $102/ton results in variations of the MESP from $0.21/gal to $1.91/gal, $0.23/gal to $1.93/gal, and $0.35/gal to $2.11/ gal for the PV, VP, and classical processes, respectively.

Ethanol production by each process requires a large amount of steam, particularly in the VFA recovery unit. Therefore, the MESP is strongly impacted by the steam price, to a greater extent even than its dependence on FCI. A ±25% change in steam price from $9.51/ ton to $15.85/ton results in MESP changes from $0.89/gal to $1.23/ gal, $0.90/gal to $1.25/gal, and $1.03/gal to $1.45/gal for the PV, VP, and classical processes, respectively. As can be seen from Figs. 5e7, the MESP shows greater sensitivity to changes in steam price for the classical process than for the PV and VP processes. This is mainly because the classical process is more energy intensive and consumes more steam than the hybrid processes. Changes in FCI

Fig. 6. Impact of changing key parameters from their base values for the VP process.

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Fig. 7. Impact of changing key parameters from their base values for the classical process.

and IRR have a similar impact on the MESP, and a ±50% change in these parameters results in a $0.23/gal change in MESP for the PV and VP processes and a $0.28/gal increase for the classical process. Changes in the digestate selling price, membrane cost, membrane life, refrigerant cost, and permeate flux have the lowest impact, with ±50% changes in base values having an impact of less than ±$0.05/gal on the MESP for the PV and VP processes. Given the high impact of the alcohol co-product selling price, it can be concluded that further refining the co-products to generate pure propanol and butanol products to sell at higher prices will have an appreciable impact on plant economics, and a further decrease in the MESP can be expected. The high impact of biomass price shows the necessity of developing enhanced artificial brown algae cultivation technology to provide higher yields and eventually cheaper biomass prices. In addition, the impact of steam price on MESP shows that the development of alternative processes for more energy efficient recovery of VFAs from fermentation broth will improve the viability of mixed alcohol production from brown algae through a VFA platform. 3.4.2. Impact of VFA yield The chemical composition of brown algae varies considerably between species and habitats throughout the year [120]. Brown algae exposed to seasonal changes accumulate mannitol and laminaran during the light season (i.e., spring to autumn) and consume these stored carbohydrates during the dark season. Differences and seasonal variations in the chemical composition of brown algae species eventually influence the VFA yield from brown algae. Based on the results of Pham et al. [80], the VFA yield can vary between ~0.307 and 0.412 g VFA/g dry L. japonica. Therefore, a single point sensitivity analysis is performed on the total yield by changing it from the base value of 0.35 to 0.30 and 0.40 g VFA/g dry brown algae. The process unit equipment are resized and capital costs are recalculated for the new flow-rates. The TCI are calculated to be $65.3 and $74.6 million for VFA yields of 0.30 and 0.40 g VFA/g dry brown algae, respectively. Fig. 8 shows the results for the impact of VFA yield on the MESP for the PV process. The MESP is $1.41/gal and

$0.87/gal for VFA yields of 0.30 and 0.40 g VFA/g dry brown algae, respectively. Results show that the VFA yield has a significant impact on the MESP. Therefore, the use of brown algae species with higher volatile-solid contents and development of high-yield

Fig. 8. Impact of VFA yield on the MESP for the PV process.

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fermentation technology can significantly enhance the plant economics. The use of modern biochemical and genetic tools to identify suitable microorganisms and genes for hydrolysis and fermentation of brown algae can play an important role in the development of high-yield fermentation technology. 3.4.3. Impact of acetic acid recovery A larger portion of VFA can be extracted from the fermentation broth by increasing the flow-rate of MTBE in the extraction column of the VFA recovery unit. However, not all of the acetic acid in the fermentation products is recoverable because part of the acetic acid remains dissolved in the aqueous phase. In addition, a higher MTBE flow-rate increases the energy and capital costs required to handle and separate MTBE from VFAs. Therefore, an economically optimal recovery value should be obtained to minimize the MESP of the process. For this purpose, a sensitivity analysis is performed on acetic acid recovery in the extraction column. The amount of acetic acid recovered is changed from its base value of 95%e90%, 85%, and 80%. A design spec was assigned to manipulate the MTBE flow-rate to reach these recovery ratios. The capital costs are recalculated for the new flow-rates. The TCI varied from $65.7 million for the base case of 95% to $62.9, $61.1, and $59.1 million for 90%, 85%, and 80% acetic acid recoveries, respectively. The annual utility costs decrease from $20.2 million/year for the base case of 95% to $19.4, $18.9, and $18.5 million/year for 90%, 85%, and 80% acetic acid recoveries, respectively. Consequently, the MESP changes from $1.062/gal for the base case of 95% to $1.044/gal, $1.047/gal, and $1.050/gal for 90%, 85%, and 80% acetic acid recoveries, respectively. The MESP is more sensitive when the acetic acid recovery changes from 95% to 90%. Later, by decreasing the amount of acetic acid recovery to 85% and 80%, the MESP rises because of the decrease in revenue from the alcohol products. The minimum MESP is obtained at 90% acetic acid recovery. Results also show that changes in acetic acid recovery ratio above 80% do not have a large impact on the MESP mainly because increased revenues from the sale of alcohol products compensate for the increased energy and capital costs for the recovery and separation of MTBE from VFAs. 3.4.4. Impact of plant scale Choosing the appropriate plant scale is an important economic factor, known as economy of scale, and can significantly influence the economics of the process. Studies on thermochemical conversion of microalgae have shown that plant scales above 100,000 ton/ year are required to make the ROI (return on investment) economically attractive [121,122]. As can be seen from Eq. (2), the capital costs of the process do not increase linearly with capacity. Therefore, larger equipment sizes enable lower cost-per-unit capacity [103]. However, larger plant scales are limited by availability and access to the biomass and might result in increased transportation costs. In addition, maintaining the optimum performance with larger equipment is also important and may not be feasible above specific scales. In this section, a sensitivity analysis is performed on plant scale for the PV process to evaluate its impact on MESP. For this purpose, the plant scale is changed from a base scale of 400,000 ton/year to 300,000, 500,000, and 600,000 ton/year. The Aspen Plus simulation is updated for the new scales. The results of the simulation are used to recalculate the capital costs of each unit. The TCI is changed from the base value of $65.7 million at a plant scale of 400,000 ton/year to $51.2, $80.0, and $92.7 million at plant scales of 300,000, 500,000, and 600,000 ton/year, respectively. Fig. 9 shows the impact of plant scale on the MESP for the PV process. MESP values of $1.09/gal, $1.05/gal, and $1.02/gal are calculated for plant scales of 300,000, 500,000, and 600,000 ton/year, respectively. The maximum MESP difference of $0.04/gal between plant scales of

Fig. 9. Impact of plant scale on MESP for the PV process.

400,000 and 600,000 ton/year shows that larger plant scales do not have a significant impact on the MESP, and a plant scale of 400,000 ton/year is a more reasonable scale considering the limitations on the availability of brown algae and its current cultivation trends.

3.5. Carbon footprint of the alcohol recovery method Based on the methodology explained in Section 2.6, the carbon footprint of the alcohol dehydration unit is calculated for the hybrid PV/distillation, hybrid VP/distillation, and classical molecular sieves/distillation processes. The contribution of each utility to total emissions is shown in Fig. 10. The total CO2-eq emissions are calculated to be 23.120, 30.190, and 62.160 kton/year for the PV, VP, and classical processes. The total emissions of the classical process are almost three times those of the PV process and twice those of the VP process. In this respect, the hybrid PV/distillation process is a cleaner alternative than the classical and VP processes. The emissions attributable to cooling are higher in the PV and VP processes than in the classical process, mainly because of the refrigeration energy that is needed to condense the permeates in the membranes. However, the CO2-eq emissions from heating in the classical process are much higher than those of the PV and VP processes; this eventually leads to the higher total CO2-eq emissions from the classical process. The CO2-eq emissions of the VP process are 30% higher than those of the PV process, mainly because of differences in the operating temperature and process structure. The VP process operates at 145
C and requires MP steam at 185
C for provision of heat; in contrast, the PV process operates at 75
C so LP steam at 164
C can meet the heating demands. Furthermore, the entire feed stream in the VP process must be in the vapor phase before contacting the membranes, which requires a large energy input. In contrast, only the permeates are evaporated in the PV process, for which heating by LP steam is sufficient. Therefore, it is expected that, even at lower VP temperatures, the CO2-eq emissions of the PV process would be less than those of the VP process.

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Acknowledgment This research was supported by Basic Science Research Program through the National Research Foundation of Korea (NRF) funded by the Ministry of Science, ICT and Future Planning (2013R1A1A1A05004852). Appendix A. Supplementary data Supplementary data related to this article can be found at http:// dx.doi.org/10.1016/j.energy.2015.10.123. References

Fig. 10. CO2-eq emissions of the alcohol recovery and separation unit for the hybrid PV/distillation, hybrid VP/distillation, and classical molecular sieve/distillation processes.

4. Conclusion In this study, the economics of mixed alcohol production from brown algae through a VFA platform were evaluated. The results showed that this is an industrially viable process, and the selling price of ethanol is comparable to its current market price. The results also showed that VFA recovery is the limiting step in terms of energy requirements and capital costs. Therefore, the development of new technologies for VFA recovery that increase the energy efficiency or VFA concentration in the fermentation broth is recommended. Three alternative methods of alcohol dehydration, recovery, and separation e hybrid PV/distillation, hybrid VP/ distillation, and classical molecular sieves/distillation e were evaluated. The results showed that the hybrid PV/distillation process is superior in terms of MESP, energy consumption, and CO2-eq emissions. Experimental studies show that permeate fluxes with the VP process can be up to four times larger than those with the PV process. However, the results also show that the energy requirements to vaporize the entire feed stream in VP can result in utility costs that are approximately 30% higher than those for PV; also, VP has higher CO2-eq emissions. Therefore, the application of the VP process is recommended for processes in which the feed stream contains solids and impurities, such as sulfates, that can affect the durability and performance of PV membranes. The sensitivity analysis identified that the selling price of the co-product alcohols, biomass price, and variation in VFA yield caused by seasonal changes in the chemical composition of brown algae or changes in the performance of the fermenting microorganisms in the fermentation reactors are the main parameters that influence the MESP. Therefore, any increase in co-product selling price or VFA yield or decrease in biomass price can strongly contribute to plant economy. To achieve this, further refinement of the heavier alcohols to produce pure butanol and propanol for sale at higher prices is recommended. High-yield artificial cultivation of brown algae, which should result in lower biomass prices, should be further developed and assessed. In addition, identification of suitable microorganisms and genes for fermentation of brown algae and genetic modification of fermenting microorganisms could lead to increased VFA yields and enhanced process economics.

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