KET050 Feasibility Studies on Industrial Plants Dept. of Chemical Engineering, Lund University

Transcription

1 KET KET050 Feasibility Studies on Industrial Plants Dept. of Chemical Engineering, Lund University A Feasibility Study on Retrofitting a starch based ethanol production to lignocellulosic materials Presented to Agroetanol Sweden May 15, 2016 Principal investigators: Joaquin Gomis, Erik Lorin, Miguel Sanchis, Erna Trokic, Jakob Ulmestig Tutors: Anders Gundberg Agroetanol Ola Wallberg and Borbala Erdei Department of Chemical Engineering Disclaimer This report was prepared as a project in the course Feasibility Studies on Industrial Plants, (KET050), Department of Chemical Engineering, Faculty of Engineering, LTH, Lund University Sweden in cooperation with the Swedish company Agroetanol. Neither Lund University nor the authors of this report or Agroetanol may be held responsible for the effects following from using the information in this report. Nor the authors, Lund university or Agroetanol makes any warranty, expressed or implied, or assumes any legal liability or responsibility for the accuracy or completeness of this information. No reproduction is authorized without the written permission from the authors, Agroetanol, or Lund University.

2

3 Table of Contents 1. PROJECT AIMS AND SCOPE BACKGROUND Biofuels and renewable energy G ethanol processes The raw material Fermentation steps: liquefaction, saccharifaction and fermentation Distillation and by-product valorization G ethanol processes The raw material. The need for a pretreatment Cofermentation vs biogas and bioethanol coproduction Integrated ethanol processes THE PROPOSED PROCESS. TENDING TOWARDS BIOREFINERIES Base case definition Pretreatment. C6 and C5 compounds separated Fermentative steps Distillation and waste valorization Process modelling Wheat straw pretreatment Ethanol production Ethanol purification Animal feed production Biogas production Biogas upgrading Results and process performance PINCH ANALYSIS. IMPROVING ENERGY CONSUMPTION Introduction. Background notes on pinch analysis Improvement targets The designed heat exchanger network... 21

4 5. ECONOMIC EVALUATION Investment cost Sizing of the units Grass roots plant cost Operating cost Fixed Capital Direct Costs Indirect Costs Profitability and sensitivity analysis CONCLUSIONS FURTHER WORK REFERENCES APPENDIX 1: CONVERSION FACTORS IN PRETREATMENT APPENDIX 2: CONVERSION FACTORS IN BIOGAS PRODUCTION APPENDIX 3: UNITS SIZES AND DIMENSIONS APPENDIX 4: EQUIPMENT CAPITAL COST ESTIMATION APPENDIX 5: OPERATING COST CALCULATIONS APPENDIX 6: PRICES FOR CONSUMABLES USED... 47

5 1. PROJECT AIMS AND SCOPE This project is envisioned as a mean of evaluation in the course Feasibility Studies on Industrial Plants (KET050) in the spring semester Thus, the aim of the work exposed in the following pages is to prove that the authors have acquired the necessary abilities to pass this course, which is imparted in Lund Tekniska Högskola (LTH). At the same time, the project is meant to provide valuable knowledge to the company Agroetanol, which has been supervising the work developed together with teachers at LTH. The objective of the project is to determine whether it is feasible or not to retrofit a starchbased ethanol production to lignocellulosic materials, in particular wheat straw. The project is then considered a feasibility study and as such it does not contain in-depth calculations corresponding to a detailed solution. In other words, the aim of the project is to determine if it is worth to develop a detailed solution and not to determine the feasibility of the process to the smallest detail. As a consequence, detailed calculations such as piping and hydraulic installation, security and control systems or heating and cooling installations are not included in the scope of this project. Aspects regarding the construction of the process are neither included. The feasibility study scope includes both technical and economic feasibility. Therefore, material and energy balances for the proposed process are solved and cost estimation is also performed afterwards, in which collaboration with the local district heating company is included in the analysis. Even if there are several process alternatives to achieve the aim of the project, only the option that is considered best by the authors, providing reasoning that justifies the choice, will be analyzed to the full extent of the scope. However, in the economic part several sensitivity analyses will be performed in order to study the effect of different variables on the profitability of the process. 1

6 2. BACKGROUND The usage of fossil fuels releases greenhouse gases, which is known to hinder heat from leaving the Earth, and could in the end lead to a climate change. According to Environmental Authorities, the transportation sector stands for one fourth of the energy consumption in Sweden, up to 90% of the fuels used are fossil [1]. The Renewable Energy Directive by the European Commission is a way of lowering the greenhouse gas emission by requiring that at least 10% of the transportation fuel is based on renewable energy [2]. The transportation sector activity is not expected to decrease, rather increase, and, in line with EU s goals of limiting the greenhouse emissions, it is therefore of great importance to replace the fossil fuels for renewable sources of energy Biofuels and renewable energy There are today alternatives to the fossil fuels, for example bioethanol, which has the advantage that it consumes and releases the same amount of CO 2 and is therefore considered to have a zero net emission of greenhouse gases [3]. However, if the crops are planted on soil not used for agriculture before, then the stored carbon in the soil will be released and the overall net emission of CO 2 would no longer be zero [4]. This applies as well to the use of fossil fuels in any other step of the production chain. To avoid an increase of the greenhouse gases due to cultivation, the European Commission has set some sustainability criteria. The criteria state that a biofuel is considered sustainable if the greenhouse savings are at least 35% [5], once all life cycle emissions are included, compared to fossil fuels. Biofuels that do not fulfill the criteria are therefore not considered sustainable by the government [5]. The world s production of bioethanol is today based on crops that contain starch and sugar, which could be used instead for food, as for example maize and wheat [4]. This has become a moral problem since a large part of the world s population is starving. It is then desirable to produce bioethanol from lignocellulosic materials since there is a great availability and it does not compete with the food production industry [4]. The bioethanol based on starch and sugar is designated as first generation bioethanol (1G) while bioethanol based on lignocellulose is regarded as second generation bioethanol (2G). 2

7 2.2. 1G ethanol processes The raw material The raw material used for first generation bioethanol, such as wheat, maize or sugarcane, mostly contains starch and sugars. The selection of raw material depends on availability: mainly wheat is used in Europe while most of the ethanol produced in Brazil is made from sugarcanes [3]. Ethanol is produced through fermenting sugars, which for the sugar-based raw material means that the sugars only need to be extracted from the raw material and can be used straight away. For the starch-based ethanol, sugar must be formed by liquefaction and saccharifaction of the starch before the fermentation, since starch cannot be directly converted to ethanol [3]. The raw materials for 1G bioethanol have in common that the sugars, both the ones that naturally exist in it and the ones formed during pretreatment, consist of 6 carbons (hexoses) Fermentation steps: liquefaction, saccharifaction and fermentation The first step in converting starch to sugar is to finely ground the grain by milling. Smaller particles are easier to penetrate, which is an advantage in the following cooking and liquefaction steps. In the cooking step, water is added to the ground particles and the temperature is set to C [3]. The water makes the particles swell, the starch loses its crystallinity and becomes instead gel-like. This process is called the gelatinization and makes it easier for enzymes to attack the carbohydrate bonds. After the cooking, the liquefaction begins. The temperature is decreased to C and the enzyme alpha-amylase is added, which breaks down the starch into shorter carbohydrate chains called dextrins [4]. The liquefaction decreases the viscosity of the gel formed from the cooking and the result is more liquid-like. The liquid is then exposed to the enzyme glucoamylases, which hydrolyzes the dextrin to glucose. This step is referred to as saccharification and is performed at maximum 50 C [3]. The main difference between the two different enzymes is that alpha-amylases break random bonds in the starch, while the glucoamylases hydrolyze the dextrin at a specific point. The formed glucose molecules are then fermented at 32 C, usually by yeast, to ethanol. After the fermentation an ethanol concentration of 10-12% is commonly achieved [3]. The cooking occurs for about one hour, the liquefaction can take between minutes while the fermentation lasts hours [3, 4]. 3

8 It is possible to run a simultaneous saccharification and fermentation (SSF). This configuration is alleged to decrease process time and, since the formed glucose is converted to ethanol, it is also alleged to decrease end-product inhibition [3]. Although there is one disadvantage: the temperature set is a compromise between the two different processes. This means that none of the two processes are carried out at their optimal temperature [3] Distillation and by-product valorization Distillation is the first step to achieve a high concentration of ethanol after the fermentation. The distillation is a very high energy consuming step and it is therefore of high importance to make it as efficient as possible. The energy demand is increased considerably when the feed to the distillation has an ethanol concentration below 4% [6], but this is not a concern for 1G ethanol production since the concentration is usually higher, as mentioned above. The mixture fed to the distillation contains water, which due to azeotropic behaviour gives a highest ethanol concentration of about 90 mol% at atmospheric pressure. This was ensured by the authors by running a binary mixture analysis with the components in Aspen Plus. To reach 100% pure ethanol further processing is needed, some alternatives are low-pressure distillation, drying the ethanol by using molecular sieves or using membranes. The stillage from the distillation contains fibre, protein and other components that have a high nutritional value [4]. The stillage is separated into a solid and a liquid fraction. The liquid fraction is evaporated before being mixed with the solids again. This mixture is then dried and the product is called distiller s dried grains with solubles (DDGS), which can be sold as feedstock for animals G ethanol processes The raw material. The need for a pretreatment Bioethanol based on lignocellulosic materials is denoted as second generation (2G) ethanol and the raw material can for example be wood, agricultural waste or lignocellulosic waste. The raw material consists of cellulose, hemicellulose and lignin, but the amount of each component can differ depending on what type of material is used [3]. Cellulose consists of glucose molecules that are linked together into chains in the cell wall, the length of the chains differs between different materials [3]. Cellulose consists of both crystalline and amorphous structures and does not dissolve easy in water [3]. Hemicellulose consists of many different types of sugars, the structure comprises a backbone formed by sugar residues and branches of sugars attached to the backbone [3]. The sugars can be both hexoses and pentoses, which type that is more common depends on the type of plant. Due to its branched structure, hemicellulose will be easier to hydrolyse than cellulose [3]. 4

9 Lignin consists of polymers that are linked together, giving support to the cell wall as well as protecting it from degradation [3]. The cellulose, hemicellulose and lignin are strongly connected to each other, which makes lignocellulosic raw material hard to degrade, since the plants are developed to withstand such stress [3]. To make the biomass more accessible for enzymes to break down into monomeric sugars, there must be some form of pretreatment that disrupts the strong structure of the raw material. There are different types of pretreatment: some break down the physical structure of the biomass while some change the chemical composition. Milling and chopping the biomass into finer size is an example of physical pretreatment, while pretreatments with acids and bases that solubilize the biomass are chemical [4]. Physiochemical pretreatment is a combination of both physical and chemical pretreatment, steam pretreatment with added acid being one example. Other pretreatment methods are ammonia fiber explosion, hot water and organosolv [4]. Depending on which pretreatment is used, different types of side-products are formed when the lignocellulosic material is degraded. These compounds are called inhibitors since they can inhibit the following enzymatic and fermentation steps by creating a toxic environment for the microorganisms or decreasing the ethanol production [4]. The type of inhibitors and their concentration depend on pretreatment method and material used. Therefore, the choice of pretreatment is important because it affects all the following steps in the process. Using a high concentration of acid gives low concentration of inhibitors since it enables solubilisation of cellulose and hemicellulose at low temperature and pressure, but the amount of acid makes it expensive. If the amount of acid is lowered, the temperature and pressure need to be increased, which results in a higher concentration of inhibitors because the released sugars are further degraded into these by-products [3] Cofermentation vs biogas and bioethanol coproduction Depending on pretreatment, the hemicellulose and lignin can end up in the liquid fraction or in the solid fraction with the cellulose, which consequently directly affects how the following process steps should be performed [4]. The following steps are enzymatic hydrolysis, which degrades the material into sugars, and fermentation, which converts the formed sugars into ethanol, much alike the 1G process. The difference between 1G and 2G process is that the latter has lignin and pentoses, mostly xylose, derived from the hemicellulose. The lignin can be used to generate heat by burning or pelletizing it while the pentoses can either be fermented to ethanol or anaerobically digested to produce biogas [4]. The amount of ethanol produced will naturally be higher if the pentoses are fermented, since there will be more sugar to be converted into ethanol, but the yeast normally used in fermentation of hexoses (Saccharomyces cerevisiae) cannot be used to ferment pentoses [4]. 5

10 There are other yeasts and bacteria that can ferment pentoses, but it is preferred to use one microorganism to ferment both the hexoses and pentoses since it would make the process simpler [3]. Thanks to genetic engineering, there are today microorganisms that are able to ferment both of the sugars, but they often have more affinity to ferment glucose than xylose. If the glucose concentration can be kept low, a higher amount of the xylose can then be fermented [3]. Still, there are no microorganism today that can achieve same yield and productivity as the S. cerevisiae [7]. There is also legislation for the usage of genetically modified products, which can cause problems if the products are aimed for human consumption [8, 9]. To avoid these disadvantages, the pentoses can, as mentioned above, be used in biogas production instead, but this will decrease the ethanol yield since sugar is removed from the ethanol production. Biogas is produced in an anaerobic digester, in which different types of microorganisms break down the material and produce methane. A higher temperature increases the rate of production, but the process still takes days and therefore a large plant is needed to be able to convert all the material [4]. The biogas can be used to supply the production plant with energy or be sold. The separation of lignin and hemicellulose can be performed at different stages in the process and the way of performing it is different whether they are in the liquid or solid phase. However, removing the hemicellulose after the pretreatment, before the enzymatic hydrolysis, can give a lower concentration of inhibitors and removing the solids gives the opportunity to recirculate enzymes and yeast [4] Integrated ethanol processes When harvesting the raw material for 1G production, 2G raw material is generated, for example wheat straw after harvesting wheat. This material was historically considered as waste and therefore burned [4]. This does not fit well with current general opinion about utilizing all of the raw material. A stand-alone 2G plant only fermenting glucose does not, in general, reach sufficient ethanol concentration for the distillation to be economical due to the large energy consumption. Integrating a 2G plant with a 1G, which usually has an ethanol concentration above 10% [3], could give a combined concentration that makes the distillation economically feasible. The integration can take place at several places in the process. However, introducing the 1G line before the fermentation can be beneficial for both of the processes, since the sugar concentration in the 1G line is often diluted anyway and the 2G has to reach a higher concentration [4]. 6

11 3. THE PROPOSED PROCESS. TENDING TOWARDS BIOREFINERIES The process that defines the core of this project s solution is designed through the methodology proposed in [10]: based on a literature research a first base case is created, then a more detailed process is synthetized and finally the design is refined through optimization and cost estimation. In this project the second step in the method is carried out in Aspen Plus, which not only allows solving the material and energy balances in the process but can also be used as a tool to optimize the design afterwards Base case definition As commented in the background section, there are two main possibilities to design an ethanol process based on lignocellulosic materials: coferment C6 and C5 sugars or coproduce ethanol and biogas. The latter option leads to higher material and energy efficiency as well as to a minimum ecological footprint [11, 12]. Moreover, some concerns could arise if compounds which have been in contact with the genetically modified organisms required in the cofermentation route end up mixed in the animal feed, since the animals will in turn be used to provide food for humans. For all these reasons it is decided to base the proposed process on the biogas coproduction route, shaping the process towards the biorefinery concept. Therefore, the main part of the C6 sugars in the raw material is used to produce bioethanol, while the majority of C5 sugars is used to produce biogas. The problem with this route is that the ethanol concentration after the fermentation step is very low, which leads to an increased energy consumption in the subsequently ethanol purification [3, 4]. To overcome this disadvantage, 1G raw material is also used in the fermentation to increase the ethanol concentration. The combination of these two principles leads to an integrated plant in which ethanol and biogas are coproduced (Figure 1). In the following sections one can find a more detailed explanation of the proposed process Pretreatment. C6 and C5 compounds separated There are several ways to pretreat lignocellulosic materials, but one of the most extended technologies over the world is steam explosion [3, 4]. This is because this is a very energy efficient technology, which leads to a lower energy consumption [13]. 7

12 Moreover, in steam pretreatment xylan is partially hydrolyzed [13], which means that most of the C5 sugars tend to be transferred to the liquid phase while most of the C6 sugars remain in the solid phase along with lignin. This is very convenient for the proposed process because with a simple filter it is then possible to separate each kind of sugars. Figure 1. Base case with operating temperatures and pressures. Therefore, due to its low energy consumption and the easiness to separate C5 sugars from C6 sugars, steam explosion is chosen as the pretreatment technology in the proposed process. The conventional steam explosion vessel is coupled with a filter unit so C5 sugars can be sent to an anaerobic digester to produce biogas while C6 sugars can be sent to the ethanol production line Fermentative steps In the proposed process there are a total of four steps in which biocatalysts are used because before the fermentation the 2G material has to be hydrolyzed and the 1G material has to be liquefied and saccharified [3]. It is possible to add more than one of the required biocatalysts in just one vessel to carry out several of these steps at the same time [3]. It is chosen to perform the saccharification and fermentation in the same vessel (SSF), but the liquefaction and hydrolysis are performed separately. This alternative will lead to a higher investment cost, but at the same time the different process can be carried out at their respective optimal temperature [3]. However, the reason why SSF was chosen is that the inhibitory effects in the reactors are lower. Thus, the material efficiency of the process is improved, which will overcome the higher investment cost in the long term. 8

13 Moreover, if the hydrolysis is performed in a separate vessel it is possible to eliminate the lignin before the fermentation, a separation process much easier than separating it once it is mixed with the protein coming from wheat in the fermenter Distillation and waste valorization The stream coming out from the fermenter contains large amounts of water and other residual components apart from the desired ethanol. It is then necessary to separate the ethanol by means of distillation [4, 3]. Although it cannot be seen in the block diagram, it is decided to use a cascade of distillation towers to retrieve the ethanol because this way it is possible to perform heat integration between the different columns, in a similar fashion as it can be seen in [4]. Since the system ethanol-water has an azeotrope, in order to reach high purity in the ethanol stream it is not enough with the distillation step [4, 3]. In this case molecular sieves are chosen to remove water once the azeotropic concentration is reached. The bottom product coming from the distillation step contains proteins and other valuable components that can be turned into animal feed [4]. However, a large amount of water is also present in the stream, which needs to be removed to sell the product as animal feed. In order to reduce the moisture of the stream a first filtration step is carried out, but a considerable amount of the desired protein remains in the liquid phase. Therefore, the liquid is sent to an evaporator to remove part of the water so that the solid coming out from the evaporator can be mixed with the solid coming from the filter. After the filtration and evaporation, the solid is sent to a drying step in which an appropriate moisture content is reached [4]. Following the design in [4], it is decided to preheat the solid before the dryer in order to reduce the energy consumption in this unit Process modelling Based on the process defined in the base case, a more detailed design is developed in Aspen Plus. The approach to create this model was to divide the process in six different sections (Figure 2) and then integrate the different models to compose the overall process. The Aspen model accounts for the technical feasibility of the process since the program solves both material and energy balances in each of the units. In this section detailed explanations about how each part of the process is modelled can be found. All the results and figures presented in this section are obtained using the NRTL model to predict the physical properties of the different components involved in the process, except for the biogas purification section in which ELECNRTL is used to account for the electrolytic behavior of the components. The chosen model was NRTL according to the choices that other authors who studied deeper the field made [4, 14, 15]. 9

14 Figure 2. Overall process divided in the six different parts treated separately during modelling. Gomis-Lorin-Sanchis-Trokic-Ulmestig 10

15 Wheat straw pretreatment The pretreatment section is composed of a reaction vessel followed by two flash vessels in which steam is partially removed. Subsequently a filter unit is used to separate the C6 sugars from the C5 sugars, so each of them can be sent to its corresponding section (Figure 3). Figure 3. Pretreatment model including temperatures, pressures and unit duties. The pretreatment reactor is modelled as an RStoic reactor operating adiabatically, with fixed conversion values. The different reactions that can occur in the reactor as well as their stoichiometry were extracted from one of the papers in [4]. The conversion factor for each reaction was calculated using the information in [16], being it possible to find the calculation method in Appendix 1. Apart from the actual reactions that take place in the reactor, the transformation of xylan and arabinan to C5-degradables is considered to close the mass balance for these components. In addition, two more reactions are also added to account for the solubilization of part of the lignin and ash (Table 1). The amount of steam introduced in the reactor is such that the duty in the reactor is zero, that is, steam is used to provide heat as well as enable the pretreatment reactions. Apart from steam, fresh water is used to dilute the stream to match the experimental results in [4], that is, 9.7% water insoluble solids (WIS) content after flashing the steam. Using these specifications the composition of the final liquid after filtration matches the experimental results in [3], which supports the validity of the model. Following the methodology in [4], the filtration step is modelled as a filter with 95% particle retention. The problem here is that in Aspen Plus is not possible to define lignin and ash as soluble components, just as solid components. Therefore the solubilized part of these components remains in the solid instead of being transferred to the liquid, as it happens in reality. 11

16 Table 1. Reaction scheme considered in the pretreatment reactor. Reaction stoichiometry Conversion factor Glucan 1 Glucan+H 2 O Glucose Glucan HMF+2H 2 O Xylan 3 Xylan+H 2 O Xylose Xylan Furfural+2H 2 O Xylan C5-deg+2H 2 O Arabinan 6 Arabinan+H 2 O Arabinose Arabinan Furfural+2H 2 O Arabinan C5-deg+2H 2 O Others 9 Lignin (s) Lignin (l) Ash (s) Ash (l) Acetate Acetic acid In order to solve this problem a separator unit is included in the model. This unit separates the solubilized fraction of lignin and ash and transfers them to the liquid phase. Combining the effect of both separator and filter it can be said that the model provides good estimations of both phases composition. Finally, after the filtration step, water is added to the solid phase to reduce the WIS content to 15%. The reason behind this dilution is that SSF yields more ethanol at WIS contents between 10-20% [4] Ethanol production The pretreated wheat straw is hydrolyzed and then sent to a fermenter to produce ethanol. In order to increase the ethanol concentration in the fermentation broth, liquefied wheat is also introduced during the fermentation (Figure 4). The amount of wheat used in the process is such that the fermentation broth contains 8 wt% ethanol. The hydrolysis reactor, fermenter and liquefaction reactor are modelled as RStoic reactors operating isothermally. The operating temperatures are 50, 32 and 90 C respectively. The reaction scheme in the hydrolysis reactor and fermenter (Table 2) is extracted from the information in one of the papers in [4]. The reaction scheme in the liquefaction reactor (Table 2) is obtained through information provided by Agroetanol. In the fermenter carbon dioxide is formed along with ethanol and other by-products. In reality CO 2 is vented off the reactor at the same time it is formed, but in Aspen Plus this is modelled through a separator unit that removes all the CO 2 after the fermentation step. 12

17 Figure 4. Ethanol production model including temperatures, pressures and unit duties. The model includes several heat exchangers to preheat or cool the streams so that they have the same temperature than the reactor they are sent to. Apart from these heat exchangers, the model includes also a filter unit to remove the lignin before the fermentation step because after the fermentation reactions have taken place it is very difficult to separate the lignin from the rest of components [3]. The unit is modelled as a filter press with 95% particle retention. Table 2. Reaction scheme considered in the hydrolysis, liquefaction and fermentation reactors. Reaction stoichiometry Conversion factor Hydrolysis 1 Glucan+H 2 O Glucose Xylan+H 2 O Xylose 0.78 Liquefaction 3 Starch+H 2 O Glucose 0.84 Fermentation 4 Glucose 2 Ethanol+2CO Glucose+2H 2 O 2 Glycerol+O Xylose+2H 2 O 2 Xylitol+O Glucose 3 Acetic acid Glucose+1.5O 2 3 Succinic acid+3h 2 O Starch+ H 2 O Glucose

18 Ethanol purification The NRTL model predicts an azeotropic point for the mixture ethanol-water at 85 mol% ethanol content approximately. Therefore, to extract the ethanol from the fermentation broth it is necessary to combine distillation with another separation unit, which in this case is a molecular sieve (Figure 5). Figure 5. Ethanol purification model including temperatures, pressures and unit duties. The distillation step is modelled as a cascade of towers operating at different pressure because with this strategy the towers at higher pressure can supply heat to the towers at lower pressure [4]. In this project it was chosen, after testing several combinations, to use a 25-tray stripper operating at 0.3 bar followed by a 45-tray rectifier operating at 2.5 bar. The main advantage with this configuration is that the first tower operates at lower temperatures than the second one. This is important because the protein present in the fermentation broth can precipitate, which will cause clogging of the trays, if the temperature is too high (around 100 C). Thus, the first tower removes the protein and other by-products present in the fermentation broth while the second tower, now free of the precipitation risk, operates at higher temperatures to remove water until the mixture is close to the azeotrope (84 mol% ethanol). The first bottom product is subsequently used to produce animal feed and the second bottom product is regarded as waste water. Once the mixture is close to the azeotropic composition, it is sent to a molecular sieve in which the rest of the water is removed. The molecular sieve is periodically washed with the bioethanol produced to regenerate it. This is modelled in Aspen Plus through a separator unit which removes the water, but also 20% of the ethanol to represent the amount lost during the washing. The waste stream is recycled to the distillation cascade to avoid losses of the desired product. 14

19 Animal feed production The bottom product from the first distillation tower contains valuable by-products, such as protein or nitrogen free extract (NFE), which can be sold as animal feed (DDGS) to improve the economy of the process. The problem is that the water content in the stream is too high so an area to dry the solids is required (Figure 6). Figure 6. Animal feed production model including temperatures, pressures and unit duties. First of all the bottom product coming from the stripper column is pressurized to atmospheric pressure and subsequently filtered, being modeled the unit as a filter press with 95% particle retention. It is assumed that half of the protein goes to the liquid phase and the other half to the solid phase, which is achieved through a separator unit following a similar principle than in the pretreatment for the solubilization of lignin and ash. Since a considerable amount of protein is lost in the liquid phase, the liquid is sent to flash unit in which water is removed through applying heat. The amount of heat applied is such that the outgoing solid has a similar dry content as the solid coming from the filter, so the two of them can be mixed. The solid is then preheated to 140 C and subsequently sent to a dryer in which an appropriate dry content is reached. In this project it was chosen to simulate the dryer through combining different basic units of Aspen Plus instead of using the drying unit because the design of the latter is more difficult. To start with, heat is applied to vaporize the water contained in the solid, which is then removed in a flash vessel. The amount of heat consumed is such that the dry matter of the outgoing solid from the flash vessel is 88 wt%, which is the same value used in [4]. The vapor removed in the flash vessel is partially (90% of the stream) recycled and superheated so it can be mixed with the incoming solid to reduce the energy consumption in the drying step. 15

20 Biogas production The liquid phase generated during the pretreatment step, which contains mainly the C5 sugars present in the wheat straw, is sent to an anaerobic reactor to produce biogas. Apart from the actual reactor, the Aspen model comprises several fictional units to determine how much organic matter is degraded to form biogas (Figure 7). Figure 7. Biogas production model including temperatures, pressures and unit duties. The reactors ALL-COD, FRACCON and TOTCON do not represent process units but calculation units that the model needs to determine the amount of biogas that can be formed from the ingoing liquid stream. The reactor ALL-COD is used to gather all the organic matter under a same component, which is necessary because the methane yields are expressed per unit of COD introduced in the reactor [4]. The other two reactors are used to calculate the fractional conversion of COD into methane. The anaerobic digestion of the ingoing liquid is modelled through the reactor CODMET, which is an RStoic reactor operated isothermally at 40 C. The model considers that part of the organic matter in the liquid is degraded to form methane and CO 2, being determined the fractional conversion through a calculator block that uses the information gathered in FRACCON and TOTCON (see Appendix 2). After the anaerobic digester the water present in the stream is removed, which is modelled through a separator unit because the biogas will directly separate from the water, since it is in a different physical state. The raw biogas is subsequently pressurized in a compressor and sent to the purification area to remove the CO Biogas upgrading In order to reach enough quality to sell the biogas as a fuel, it is necessary to remove the CO 2 present in the gas. This removal is achieved through a reactive absorption in which an amine solution is used as the solvent. A regeneration tower is also considered in the model to minimize solvent consumption (Figure 8). 16

21 Figure 8. Biogas purification model including temperatures, pressures and unit duties. The first step in the purification process is to contact the raw biogas with an amine solution for the CO 2 to be absorbed by the liquid, a process that is enhanced by the reaction between the amine and the CO 2. The outgoing gas is then the desired product: a gas which is mostly methane and therefore can be sold as a fuel. The tower is modelled as an 8-tray absorption tower and the solvent used is a 30% MDEA solution. The outgoing liquid is sent to a stripper tower operating at lower pressure (1 bar instead of 3 bar) in order to remove the CO 2 absorbed so the liquid can be used again in the absorption tower. The stripper is modelled as a 7-tray distillation column and it allows removing almost all the CO 2 absorbed from the raw biogas. In both absorption and regeneration tower there are solvent losses, which need to be compensated for the recycling loop to work. Therefore, there is a reposition of solvent before the regenerated liquid is sent back to the absorption tower again. The size of the reposition stream is determined through a calculator block that calculates the MDEA lost and a design specification that makes sure that the total flow of solvent is constant Results and process performance The proposed process consumes ton/year of raw material ( of straw and of wheat) and produces m 3 /year of ethanol. This means that m 3 of ethanol per ton of raw material is produced, which is close to the m 3 ethanol/ton of raw material achieved in the process proposed in [4] and to the m 3 ethanol/ton of raw material achieved in [17]. The fact that these values are similar reassures the validity of the model. 17

22 In the process not only ethanol is produced, but also other valuable products: m 3 biogas/ton raw material, m 3 CO 2 /ton raw material, ton lignin/ton raw material and ton animal feed/ton raw material. The CO 2 production is slightly higher than the m 3 /ton raw material achieved in [4], since the purification of biogas allows recovering the carbon dioxide generated in the anaerobic fermenter. On the other hand, the amount of low-value fuel produced (lignin) is lower than the ton/ton raw material obtained in [4]. However this is compensated by the fact that animal feed is also produced in the proposed process, which is not the case in the combined plant designed in [4]. Following the working principles of a biorefinery, the use of raw material is optimized by producing several kinds of products and for this reason the only residue of the process comes in the form of waste water. In the way to obtain all the mentioned products m 3 /year of waste water is generated. Even though the volume generated is very big, it has to be noted that the waste water is only slightly contaminated, being the mass fraction of water above 98.5 wt% in all the streams. Moreover, the steam streams generated will presumably be used to supply heat to the process itself which means that their condensates will add m 3 /year to this waste generation. Regarding the energy efficiency of the process, it requires MW of heating and MW of cooling. Apart from these requirements, it has to be taken into account the energy input that the high pressure steam in the pretreatment represents. This input accounts for 4.82 MW more. Considering the energy value of the goods produced, the process produces MW in the form of ethanol, 2.76 MW in the form of biogas and MW in the form of solid fuel. This implies that the process can supply MW to the energy system while it consumes MW. It is expected for the process to consume more energy than it supplies because, along with energy products, other kind of products such as animal feed and CO 2 are produced, which is an energy consuming production. However, it is not possible to say that the process is energy efficient with such a high difference between the energy consumed and produced. In a few words, it can be said that the process has a good performance from a material point of view but from an energetic point of view the situation is not so good. Therefore, the optimization of the proposed process should be focused on improving energy efficiency rather than increasing yields and material efficiency in the chemical transformations taking place in the process. 18

23 4. PINCH ANALYSIS. IMPROVING ENERGY CONSUMPTION The performance of the process was much more inefficient from an energy point of view than from a material point of view and this is the reason why the refinement of the proposed solution comprises heat exchanges within the process to reduce energy consumption. This optimization of the process, explained in detail in the following section, is carried out in Aspen Energy Analyzer Introduction. Background notes on pinch analysis This introductory section is meant to enable the reader to understand the general concept of pinch analysis so that the results and discussion presented can be well understood and put into context. If needed, a deeper explanation of the pinch methodology can be found in [18]. In a process there are usually hot streams that need to be cooled and cold streams that need to be heated. It is therefore possible to exchange heat between these streams to reduce the need for external thermal loads in the process. Pinch analysis is a method to design this heat exchanges in an optimal way, that is, to obtain an external requirement of energy as low as possible. Based on the thermal loads and the temperature levels of the process, it is possible to determine the pinch temperature of the process. The pinch temperature is that at which the accumulated thermal load is zero, which means that the sum of the heating loads is equal to the sum of the cooling loads (taking also into account external loads). Pinch temperature is crucial in the design because the heat exchangers must follow three principles: no heat exchanger crossing the pinch, no external cooling above the pinch and no external heating below the pinch. Any heat exchanger that violates one or more of these principles will increase the external load required in the process. Keeping in mind the three ruling principles in pinch, it is possible to design a heat exchanger network to reduce the energy consumed in the process. In order to do this, it is also necessary to select a minimum temperature difference in the heat exchangers Improvement targets The stream data can be directly imported in Aspen Energy Analyzer from the Aspen Plus file. These data are displayed in the form of composite curves, which represent the heating and cooling loading at each temperature level (Figure 9). It is also possible to display the grand composite curve (GCC) of the process (Figure 9), which is the addition of the two composite curves. This last graph is very important because it allows checking the pinch temperature, which in this case results to be 96.1 C. 19

24 Figure 9. Composite curves and GCC for the proposed process. Before starting with the design of the heat exchangers, it is necessary to select a minimum temperature difference (ΔT min ). The choice of ΔT min is not trivial because a large ΔT min reduces the area in the heat exchangers and thus the capital cost while a small ΔT min reduces the operating cost since more heat can be exchanged in the heat exchangers. The best value for ΔT min is the one that represents the best trade-off between capital and operating cost. In the process the lowest total cost is obtained at ΔT min =4 C (Figure 10), which seems a value too low compared to the 13.5 C proposed initially by the program. Since the cost does not increase significantly when ΔT min is increased to C (Figure 10), it is decided to design the heat exchanger network considering ΔT min =12.5 C. Figure 10. Effect of ΔT min on the total cost of the heat exchanger network. Selecting this ΔT min to design the heat exchanger network implies that the minimum heat load that can be achieved is MW and the minimum cooling load, 9.89 MW. In order to achieve these targets at least 27 shell-and-tube heat exchangers are needed, which represents a total heat exchanger area of 6937 m 2. 20

25 4.3. The designed heat exchanger network Once the potential reduction in energy consumption was determined, a heat exchanger network was designed to exchange heat within the process (Figure 11). Three external utilities were considered to provide the external energy needed in the process: cooling water, high pressure steam and low pressure steam. Figure 11. Scheme of the designed heat exchanger network. It was impossible to reach the targets because doing so required using heat exchangers which had a very small heat load (less than 0.1 MW). This means that the designed network aims at coupling the biggest loads in the process, that is, it aims at reducing the biggest contributions to the energy consumption. As a result, some of the pinch principles were infringed and because of that the external energy needed increased, making it impossible to achieve the improvement targets. Even though the proposed network leads to a heating requirement 10.2% higher than the target and to a cooling requirement 23.4% higher than the target, the total cost of the designed solution is lower than that of the targets (Table 3). This reassures the validity of omitting the heat exchangers with very small loads, since the extra capital cost is higher than the savings associated to the reduction in energy consumption. Table 3. Comparison between the designed network, the initial simulation and the targets. 21

26 In other words, the targets proposed by the program represent the thermodynamic optimum (as small external loads as possible) while the designed network represents the economic optimum (as small total cost as possible). Although the targets were not reached, the designed network allows reducing energy consumption in the process dramatically. By implementing the proposed heat exchangers, the heating requirement of the process is 52% of the original one and the cooling requirement, 26% of the original one (Table 3). This means that the process including the heat exchangers consumes less than half of the energy consumed by the original solution (without considering the 20 bar steam in the pretreatment). This energy reduction improves considerably the energy efficiency of the process, since the process provides MW in the form of different energy products while consuming only MW. Now it can be said that the energy efficiency of the process is acceptable because, even though not only energy products are obtained, more energy is delivered to the energy system than it is taken from it. As final remark, it has to be noted the profitability of recompressing the steam flashed at 1 bar after the pretreatment. Even though 0.16 MW of electricity are consumed (63.9% of the total electricity consumption), 1 MW of thermal energy at high temperature are obtained, which in turn can be used to supply heat to the second distillation tower and make it possible to run both towers without any external heat supply. This difference is stressed by the fact that electricity is much cheaper than thermal energy in Sweden [4] and because not recompressing the steam implies that only part of the load can be allocated within other process streams, which means that the cooling requirement would have been increased in addition to the increase in the heating requirement. 22

27 5. ECONOMIC EVALUATION In the previous sections it has been shown that the process is technically feasible: a variety of products is obtained with high material efficiency and also acceptable energy efficiency. In this section the economic feasibility of the process will be evaluated. To do so the cost of the process is divided into two parts: the investment cost and the operating cost Investment cost In order to get the investment cost of each unit of the process it is necessary to size them previously. Then, depending on the substance to be treated, a suitable material of construction has to be selected. Finally, the investment cost is estimated based on a cost database Sizing of the units The process units can be divided into four big groups: reactors, separation units, heat exchangers and auxiliary equipment. Each of them has to be sized in a different way. In this section it is explained how each unit is sized and the assumptions that are made, but the final results of the sizing can be found in Appendix 3. The reactors have been sized according to their residence time. From the data in Aspen Plus it is possible to get the flows going through the reactors. Multiplying the flow by the residence time it is possible to get the reactor volume, considering that the liquid volume is 70% of the total reactor volume. The residence times as well as the calculated reactor volumes are included in Appendix 3. In the separation units different ways are used to size each unit. The cost of the distillation column as well as the absorber and the stripper is estimated based on their height, diameter and number of trays [19]. This data is directly gotten from Aspen Plus. Regarding the flash vessels, it is necessary to get the diameter and the height in order to estimate the cost [19]. To do so, it is assumed that the liquid volume is half the vessel volume and the liquid residence time is 5 min [20]. With this information the vessel volume can be estimated. Then, the maximum vapor velocity is calculated using Souders-Brown equation [20] and the minimum cross section area, and thus the minimum diameter, is determined. For the evaporator what it is needed is the area [19], which can be obtained from Aspen Plus. The filters are assumed to be belt filter press, which is the same type used in [4, 17]. The residence time is assumed to be 45 min and the cake thickness 32 mm [21]. With the residence time and the flow rates the filter volume is obtained, based on a filter press volume calculator [22]. Then, the filter area is calculated by dividing the volume by the thickness. 23

28 In order to size the molecular sieve column, it is assumed that the residence time is 10 min, the pellets volume is half the vessel volume and the diameter/height ratio is 11:1, which are reasonable values for a molecular sieve column [23]. This way the diameter and the height can be obtained. Regarding the heat exchangers, it is possible to obtain the area from Aspen Plus, which is what is needed to get the heat exchangers cost [19]. With respect to the auxiliary equipment, the pumps and the compressors are the only equipment that is included, since other units such as valves or pipelines are considered negligible in comparison to the other larger units. For the cost estimation of both the pumps and the compressors it is required to know the power consumption [19], which is obtained directly from Aspen Plus Grass roots plant cost Once all the units are sized, their capital cost can be estimated. The total equipment cost of a unit is the result of the following equation. C BM = C p F BM (1) Where C p is the purchased equipment cost (at one bar and normal temperature and carbon steel as a material); F BM is the module factor, which depends on the pressure, the temperature and the material (making the equipment more expensive for higher pressures and temperatures and more expensive materials); and C BM is the bare-module cost [19]. The materials has been chosen according to the temperature and the medium based on the information given in [19], always trying to choose the cheapest option. In some cases the material choice has been limited by the availability of materials that there was in the cost database. The cost database that has been used for the calculation of the bare-module costs is the one that is given in [19]. However it was decided to use the web database in [24], which allows a much quicker search of the cost data. It has to be remarked that this database has exactly the same information as the database in [19]. After getting the bare-module cost for all the units, they have to be converted to the value that corresponds to the present time. This is done by using the Chemical Engineering Plant Cost Index (I CE ). For the values given in [24] this value is 400, corresponding to January The I CE value for May 2015 is [25]. This values are given in American dollars, and they have to converted to SEK, by using the exchange rate, which for May 2015 it was $/SEK [26]. Then it is possible to get the total direct plant cost updated to May 2015 in SEK by using the next equation [19]. G direct = I CE,2015 I CE, C f BM,i $/SEK i (2) 24

29 Once the direct plant cost is obtained, the indirect plant cost (corresponding to freight and insurances, engineering, installation overhead, fees and contingency) is the 18% of the direct plant cost [19]. Finally the auxiliary facilities (buildings and processes) are estimated as the 30% of the direct and indirect cost [19]. The Grass Roots plant cost or investment cost is the sum of the direct plant cost, indirect plant cost and auxiliary facilities. All the equipment cost for each unit, as well as the selected material, can be found in Appendix 4 and below a summary of the investment cost estimation is presented in Table 4 and Table 5. Table 4. Summary of the direct plant cost. PROCESS SECTION EQUIPMENT COST (million SEK) REACTORS SEPARATION UNITS HEAT EXCHANGERS AUXILIARY EQUIPMENT TOTAL DIRECT COST Table 5. Summary of the total investment cost. It can be seen in Figure 12 the majority of the direct cost originates from the separation units (49%) and the reactor units (41%). This is to be expected as separation units are always a major cost in chemical plants, while the large quantities of water and long residence times necessary increases the reactor sizes. Figure 12. The percentage of the direct plant cost contributed by each process section. 25

30 The major contributing separation units are the evaporators and the filter presses, together accounting for 88% of the cost for the separation units. The majority of the cost related to the reactors originates from the SSF reactor and the straw hydrolysis reactor, which together adds up to 80% of the reactor cost. This is reasonable considering that the streams have been diluted at these point, and there are significant residence times. Even though the biogas digester is the largest reactor the cost is significantly reduced as it can be constructed from concrete, which is a cheaper material. The grass roots plant cost is million SEK, which corresponds to SEK/m 3 ethanol in a year or 2475 SEK/ton of raw material in a year. These values are similar to those in [17] (15039 SEK/m 3 ethanol and 2322 SEK/ton of raw material, respectively) which supports the accuracy of the estimation and the efficiency of the designed process Operating cost The operating costs are estimated using rules of thumb. The cost for the consumables (i.e. feedstocks, solvents, catalysts & utilities) are calculated whilst most other costs are estimated as added fractions on the capital and consumables costs. In order to complete these calculations, a number of assumptions are made. These are presented in Table 6 below. All the equations used are presented in Appendix 5. Table 6. General assumptions used in the operating cost calculations. ASSUMPTION VALUE UNIT INTERNAL INTEREST RATE 10 % LIFETIME OF FACILITIES 20 Years PROCESS HOURS PER ANNUM 8000 Hours/Year Fixed Capital The fixed capital costs includes the storage of the feedstocks, products and the spare parts for the plant. As products and feedstocks generally have to be supplied batch-wise rather than continuously the cost for storage has to be taken into account. The calculation method is presented in Appendix 5, and is based on the storage time, the consumption or production, the price of the material, the interest and the lifetime of the facilities. The storage time of the feedstocks wheat straw and wheat grain are estimated to be 30 days, whereas the storage time of the products is estimated to be 30 days, except for the lignin. The lignin is estimated to have a storage time of 1 day, as it is sold to the nearby CHP plant for fuel. 26

31 The cost for the spare parts are estimated from the interest, lifetime, the maintenance cost (calculated in Direct Costs) and a factor between 10%-20% [28], depending on the factory. Due to the lack of sensitive process equipment the factor was estimated to be 10%. The determined costs are presented in Table 7 below. Table 7. Costs associated with fixed capital. FIXED CAPITAL COST (million SEK per year) STORAGE OF FEEDSTOCKS 2.45 STORAGE OF PRODUCTS 2.69 SPARE PARTS FOR PLANT Direct Costs The direct costs are the cost for the consumables and the license fees, maintenance and repair, factory staff and laboratory work. The cost for the consumables (including utilities and waste streams) and license fees are calculated from the consumptions obtained from the simulation in Aspen Plus and market prices. The prices used are presented in Appendix 6. The annual maintenance and repair cost is estimated as a fraction of the grass roots capital. The rule of thumb states that this factor should be between 2% to 10% [28]. The cost for maintenance is estimated to be around 5%. As the factory produces bulk chemicals the annual process hours is estimated to be In order to achieve this a five-shift schedule is assumed. The need for operators is estimated to be low, at 2 operators per shift. The salary is estimated to be the average for a CNC-operator in Sweden, i.e SEK per month [29]. The costs for supervisors and laboratory are estimated to be low, and are therefore both set to 10%, each from a range of 10%-20% [28]. As the production line is meant to replace a previous one, the cost for purchasing or leasing land is believed to be zero. The derived direct costs are presented in Table 8 below. Table 8. Estimated direct costs. DIRECT COST COST (million SEK per year) CONSUMABLES MAINTENANCE & REPAIR OPERATORS 2.82 SUPERVISORS LABORATORY WORK

32 Indirect Costs The indirect costs include overhear for staff, administration, distribution and sales costs and the costs for research and development. The rule of thumb dictates that the overhead should be 70% on shift personnel and 50% on day personnel [28]. It also estimates the administration costs to be 25% of the staff overhead [28]. As the price for the produced goods does not include any costs for distribution or sale this has to be included as a factor. The goods produced are, however, bulk chemicals and the sales cost are therefore assumed to be low. The distribution and sales are therefore assumed to cost 5% of the sales revenue, from a range of 0%-10% [28]. As the company mainly focuses on production the research & development costs are set to 0%. The indirect costs are shown in Table 9 below. Table 9. Estimated indirect costs. INDIRECT COST COST (million SEK per year) OVERHEAD FOR STAFF 2.26 ADMINISTRATION 0.56 DISTRIBUTION & SALES RESEARCH & DEVELOPMENT Profitability and sensitivity analysis In order to evaluate the performance of the plant the calculated operating costs have to be compared to the revenue. The net revenue is presented in Table 10 below: Table 10. Annual operating costs and revenue for the plant. 28 COST (million SEK per year) TOTAL OPERATING COSTS TOTAL SALES REVENUE NET INCOME PER YEAR As it can be seen in Figure 13, ethanol represents the majority of the revenue which can be expected from an ethanol production process. However, there are significant incomes from the sales of DDGS which means that the animal feed is also a relevant product in the process. As the annual total operating costs exceed the annual total sales income the plant cannot be profitable. In order to calculate the true profitability of the facility over the length of its life the capital cost and the interest rate has to be taken into account. The calculated net present value is MSEK which implies that ethanol should be sold at least at SEK/L (the current market price is 3.45 SEK/L).

33 Figure 13. Revenue distribution between the different products of the process. However, there is one factor that was not considered in the solution which can improve the situation: water recirculation. Recycling waste water to reduce fresh water consumption decreases proportionally the operating cost, but it has to be taken into account that it increases the inhibitors concentration in the process which may decrease its material efficiency. In order to avoid an exceedingly increase in the inhibitors concentration, the authors consider reasonable to recirculate only the two waste water streams with the lowest inhibitor concentration: water coming from the evaporator and bottom product of the second distillation tower. Recycling these two streams implies that 55% of the water consumption can be supplied internally by the process itself. Taking into account the effect of recycling water, the process losses are reduced to MSEK each year. This represents only a 1.2% improvement on profitability with respect to the original situation (minimum selling ethanol price: SEK/L), which means that, apart from polishing the solution, improvements in the process design must be done for the process to be feasible. One way to improve the process is to pay attention to the filtration step in which lignin is removed. In order to remove low molecular weight lignin ph is lowered, but during the experimental trials from which this information was derived the process was not optimized [3]. Thus, it is possible to improve the economics of the process by changing the ph at which this separation happens, although the effect becomes less relevant when the ph gets closer to the ph at which the hydrolyzation takes place (Figure 14). Increasing the ph to 3 instead of 2 in the lignin filtration has a huge impact on the process economics, the losses being reduced to MSEK/year. It is worth to investigate whether enough of the lignin precipitates at this higher ph, since a 41.8% reduction in annual losses with respect to the original design can be reached. 29