Bio-Ethanol Productions Plant in Ontario

DESIGN OF MULTI-FEEDSTOCK BIO-ETHANOL PLANT IN ONTARIO

Winter Term 2015 Department of Chemical Engineering McMaster University

By Team: Vytautas Stasiulevicius, Fahd Ilyas, Carlo Bantug, Danish Fahzal, Leo (Xiau) Zhou

A Project Report CHE 4W4 – 4W4 – Chemical Chemical Plant Design and Simulation

4W4 2015 – Syngas Fermentation

Executive Summary This document explores the feasibility of building a multi-feedstock biofuels production plant in Ontario to produce Ethanol. A basis of producing 100 million litres of Ethanol was selected for designing and costing the production plant, constrained mainly to meeting the quality standards outlined in the the Canadian General Standards Board. The production of Ethanol through through a thermochemical pathway and a biochemical pathway were studied as technological alternatives. The thermochemical pathway proceeds via Gasification followed by Fischer-Tropsch synthesis but was found to be very energy intensive and required a specific gas composition for production. The alternative, biochemical pathway involves enzymatic hydrolysis followed by fermentation, but high enzyme costs, cost-intensive pre-treatment, and low feedstock flexibility deterred the selection of this process. Instead, a hybrid production pathway was selected, referred to as syngas fermentation, which combines aspects from from both the thermochemical and biochemical processes. processes. Compared to the alternatives, the hybrid process was selected mainly due to high feedstock and gas composition flexibilities, allowing nearly any lignocellulosic material to be converted into Ethanol. Syngas fermentation was also advantageous over alternative processes due to high Ethanol yield, selectivity, selectivity, and high resistance resistance to contaminants. The feedstock of interest interest is first crushed and dried before it is sent sent to a fluidized bed gasifier to produce produce a syngas mixture. The syngas mixture goes through various cleaning and cooling stages to remove impurities before being fed into a fermenter containing a specific bacteria (clostridium Ijungdahlii) acting as a biocatalyst. The bacteria continuously converts syngas to ethanol within the fermenter, while the broth is continuously extracted and sent to distillation distillation columns to separate out the desired Ethanol to be used for fuel. The proposed plant would require an approximate $90 million investment for capital costs to establish the plant infrastructure, and would cost r oughly $110 million per year to operate and maintain. Though due to high feedstock costs costs incurred from growing the the feedstock the plant would run a net negative NPV over a 25 year project lifetime unless government subsidies were provided on the price of ethanol. The overall production process produces approximately 2 kg of CO2 equivalent emissions, emissions, comparable to ~20 ~20 kg for crude oil processes. Though it must must be noted that the final Ethanol will likely be blended with gasoline, so final emission reductions will be in the order of 6-8%, which is a great great improvement.

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Table of Contents Executive Summary ........................................................................................................... 2 Table of Contents ............................................................................................................... 3 Table of Figures ............................................................................................................. 5 Table of Tables .............................................................................................................. 6 Table of Reacions ........................................................................................................... 6 Table of Equations .......................................................................................................... 6 1. Project basis................................................................................................................... 7 1.1 Summary ................................................................................................................. 7 1.2 Economic Approximation of Process ........................................................................... 8 1.3 Relevant standards .................................................................................................... 9 2. Existing solutions ......................................................................................................... 10 2.1 1st generation feedstock ........................................................................................... 10 2.2 2nd generation feedstock .......................................................................................... 11 3. Design alternatives ....................................................................................................... 14 3.1 Hybrid process ........................................................................................................ 15 3.2 Biochemical process ................................................................................................ 16 3.3 Thermochemical process .......................................................................................... 18 3.4 Bioethanol Location ................................................................................................ 20 4. Overview of Proposed Process Design ............................................................................ 22 4.1 Process Summary .................................................................................................... 22 4.1.1 Pre-Treatment ................................................................................................... 23 4.1.2 Gasification ...................................................................................................... 24 4.1.3 Gas Cleaning .................................................................................................... 26 4.1.4 Fermentation .................................................................................................... 29 4.1.5 Distillation ....................................................................................................... 32 4.2 Design Basis ........................................................................................................... 33 4.3 Product Specifications ............................................................................................. 33 5. Process Behaviour ........................................................................................................ ........................................................................................................ 33 5.1 Normal operation .................................................................................................... 33 5.2 Start-up and shutdown ............................................................................................. 36

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5.2.1 Start-up ............................................................................................................ 36 5.2.2 Shutdown ......................................................................................................... 38 5.3 Emergency Procedures ............................................................................................. 39 6. Overall Material and Energy Balances ............................................................................ 40 40 6.1 Overall Material Balance: ......................................................................................... 40 6.1.1 Process Side ..................................................................................................... 41

6.1.2 Utilities side ................................................................................................... 41 6.2 Overall Energy Balance ........................................................................................... 43 6.2.1 Process Side ..................................................................................................... 43 6.2.2 Utility Side ....................................................................................................... 43 6.3 Stream and Equipment tables .................................................................................... 45 6.3.1 Process Side ..................................................................................................... 45 6.3.2 Utilities Side .................................................................................................... 55 7. Process Control ............................................................................................................ 58

7.1 Control Overview .................................................................................................... 58 7.2 Preliminary P&ID of Process .................................................................................... 73 8. Equipment design, sizing and costing – costing – process process side ......................................................... 76 8.1 Costing overview .................................................................................................... 76 8.2 Capital costs ........................................................................................................... 77 8.3 Operating costs ....................................................................................................... 82 8.4 NPV ...................................................................................................................... 86 8.5 Sensitivity Analysis ................................................................................................. 88 8.6 Equipment Sizing .................................................................................................... 91 Heat exchanger design ............................................................................................... 93 9. Environmental Impact ................................................................................................... 95 9.1 LCA ...................................................................................................................... 95 9.2 GHG Emissions ...................................................................................................... 97 10. Process safety ........................................................................................................... 100 10.1 Hazardous Materials ............................................................................................ 100 10.2 Process Hazards .................................................................................................. 103 11. Risk Assessment ....................................................................................................... 104

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11. 1 Technical ........................................................................................................... 104 11.2 Societal .............................................................................................................. 105 11.3 Economical ......................................................................................................... 106 Appendices ................................................................................................................... 108 Appendix 1 – 1 – Various Various Lists Relating to Process ................................................................. 108 List of Materials ......................................................................................................... 108 List of Equipment ....................................................................................................... 108 List of Symbols .......................................................................................................... 108 Appendix 2- Detailed Equipment List ............................................................................... 110 Appendix 3 - HAZOP Study ............................................................................................ 111 References .................................................................................................................... 115

Table of Figures Figure 1. A hydrolysis-based cellulosic ethanol production process Figure 2. Length of Growing Season in Ontario Figure 3. Block flow diagram of syngas fermentation process Figure 4. Typical fluidized bed gasifier configuration Figure 5. Typical biomass feeding system for fluidized bed gasifier Figure 6. Typical wet scrubber configuration Figure 7. Wood-Ljungdahlii biochemical pathway Figure 8. Typical stirred-tank bioreactor configuration Figure 9. Ratio control loop design for steam to feed ratio Figure 10. Ratio control design for Feed to reboiler utility ratio Figure 11. Ratio control structure between distillate and reflux Figure 12. Pressure control inside the gasifier unit Figure 13. Pressure control loop design for the distillation columns Figure 14. Level control structure for the fermenter Figure 15. Level control for reflux drum Figure 16. Cascaded temperature control design around the condenser E-104 Figure 17. pH control loop structure for the fermenter. Figure 18. Ratio control structure between purge stream and recycle stream Figure 19. pH control design for unit S-101 Figure 20. Pre-treatment section Figure 21. Gasification section Figure 22. Gas cleaning section of the P&ID Figure 23. Summary of Fermentation section of P&ID Figure 24. Summary of Distillation section of the process Figure 25. NPV Analysis Figure 26. Sensitivity Analysis

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13 21 22 25 26 28 30 31 59 61 62 64 65 66 67 69 71 72 73 75 75 76 77 77 88 90


Table of Tables Table 1. Economic analysis of a cellulosic ethanol plant using the biochemical process Table 2. Economic analysis of a cellulosic ethanol plant using the thermochemical process Table 3. Process side material inflows and outflows of the bioethanol plant Table 4. Utilities side overall material inflows of the bioethanol plant Table 5. Utilities side overall material outflows of the bioethanol plant Table 6. Overall energy balance of the process streams Table 7. Overall energy inflow of the utility streams Table 8. Overall energy inflow of the utility streams Table 9. Stream table of process streams entering and exiting the pre-treatment section Table 10. Material and energy inflow and outflow to the equipment of the pre-treatment Table 11. Stream table of process streams entering and exiting gasification section Table 12. Material and energy inflow and outflow to the equipment of the gasification section Table 13. Stream table of process streams entering and exiting the gas cleaning section Table 14. Material and energy inflow and outflow to the equipment of the gasification section Table 15. Stream table of process streams entering and exiting the fermentation section Table 16. Material and energy inflow and outflow to the equipment of the fermentation section Table 17. Stream table of process streams entering and exiting the separation section Table 18. Material and energy inflow and outflow to the equipment of the separation section Table 19. Stream table of utilities streams entering and exiting the pre-treatment section Table 20. Utilities side material and energy inflow and outflow to equipment of the gasifier Table 21. Sum of capital costs for each type of unit and total capital cost. Table 22. All operating costs for the syngas fermentation plant. Table 23. Cradle gate GHG emissions of ethanol produced. Table 24. List of hazardous chemicals used and produced in the biochemical plant

18 20 41 42 43 44 45 46 46 47 48 49 50 51 52 53 54 55 56 57 81 84 99 101

Table of Reacions Reactions 1-3. Reactions taking place within gasifier Reactions 4-10. Primary reactions that occur within fermenter, dependant on H2/CO content Reaction 11. Regeneration reaction of adsorbent bed.

25 30 34

Table of Equations Equation 1.Price of E100 Equation 2. Incentive calculation Equation 3. Heat transfer area Equation 4. Mean Temperature Difference Equation 5. Log-Mean Temperature Difference

86 88 91 91 91

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1. Project basis 1.1 Summary The inevitable paradigm shift away from fossil-based fuels and products within the next few decades has necessitated the introduction of alternative fuels and methods of producing power. Recent volatility in the oil market, general uncertainty in the future outlook of fossil fuels, rising oil prices and unavoidable concerns for global warming, with greenhouse gases and volatile organic compounds (VOCs) being released into the atmosphere, are factors that drive this paradigm shift. Bioethanol or ethanol is one fuel that has been researched in depth for the last decade and is a promising fuel because of several key advantages. It presents an alternative that has many similarities to fossil fuels, especially in terms of the infrastructure and supply chain, but is different in the categories that make fossil fuels undesirable, such as oil drilling and byproducts. Currently, the projected potential demand of ethanol in Canada by 2022 is 2 billion liters and the production capacity is 1.2 billion liters (United Nations, 2009). This 0.8 billion liter difference is a key financial incentive and makes ethanol a viable fuel to pursue in terms of research, development and finally implementation. In addition, governments like Canada’s that subsidize ethanol or biofuels in general give another financial incentive to ethanol fuel startups. Although ethanol is a viable alternative, in order for it to compete with the oil market, its use needs to be constrained in several ways to maintain quality, production, and profit. In the case of quality, the ethanol that is produced from an ethanol plant needs to be very pure (> 95% purity), with little to no water content and very small traces of other byproducts from upstream such as ammonia, hydrogen sulfide, methane and acetic acid. Although acetic acid is considered a byproduct here, there are numerous uses for it and it can be sold instead of discarded. This may require more investment towards separation of byproducts but presents a financial incentive to pursue production of ethanol by means in which acetic acid is also produced. Canadian regulations on gasoline supply require that ethanol be 5% v/v (volume percent) of the gasoline mixture. Although there are no explicit environmental regulations on ethanol fuel production and distribution, an 7


estimated 1 Mt CO2 equivalents per year reduction in greenhouse gas (GHG) emissions is expected on top of existing projected reductions that result from other regulations in Canada (National Resources Canada, 2013). In terms of plant and feedstock feasibility, water requirements are the most important constraints as the plant would not operate without water and the feedstock would not grow without sufficient water. The location of the facility can be closer to water sources such as lakes or ponds to meet these requirements relatively easily and without significant capital investment. Since the feedstock will be grown on arable farmland, irrigation systems would be necessary and would provide sufficient water. However, this limits the feedstock location to southern Ontario as there is little rainfall in Northern Ontario, with icy conditions and heavy snow hindering the growth of feedstock. There is also the consideration of feedstock availability in Ontario. The feedstock that will be used in this project is Miscanthus. This feedstock does not grow naturally in Ontario, which means that there is no feedstock available currently. However, conditions in southern or southwestern Ontario are highly favorable for the growth of this feedstock. Introduction of this feedstock in southern Ontario is therefore not expected to be hindered by adverse climate or weather effects, and the salinity and sand percentage of topsoil in Ontario should be alright for its growth. The exact amount of feedstock that needs to be purchased is based on 100 million liter output of ethanol per year. Another raw material that is needed to run the process apart from the water and the feedstock is the bacteria required for the fermentation process where ethanol is produced. This primarily comes from chicken yard waste. Power requirements are minimal in this plant as heaters, compressors or a large amount of pumps are not required. Energy requirements to run the unit operations are mostly fulfilled by pressurized steam.

1.2 Economic Approximation of Process The production cost of ethanol is estimated based on both fixed costs and variable costs. The fixed costs include installation, labor, maintenance and interest on investment. The variable costs include feedstock, enzyme production, utilities and waste management. 8


The following cost estimation is based on producing 20 million litres of ethanol per year. The cost of miscanthus, including planting and harvesting of the crop is estimated to be about 63-74 $/tDM (Roy, 2014). For every 1 kg of dry miscanthus, an estimated 0.360.39 litres of ethanol are being produced (Roy, 2014). Therefore in order to produce 20 million litres of ethanol, we need 55 million kg of miscanthus on a dry basis. The cost of buying 55 million kg of miscanthus on a dry basis is 4 million. Ethanol processing plant construction cost is estimated to be 45 million (Roy, 2014). Selling price of ethanol used is $3.85 per litre (U.S. Department of Energy, 2015). However this price is subjected to change depending on the demand of ethanol in the market and other competitors currently producing ethanol. The total revenue that will be generated from selling 20 million litres of ethanol is approximated at $77 million.

1.3 Relevant standards Environmental standards by the government are not used to govern the production of biofuels as they are an alternative source of energy and not as detrimental to the environment as fossil fuels. However, the environment including surrounding ecosystems and bodies of water near either the feedstock location(s) or the plant location need to be cared for. Waste gases and tailings/byproduct ponds cannot be close to the habitats of wildlife, and the surrounding ecosystem should not be greatly transformed in order to introduce feedstock or to build a chemical plant. Manufacturing standards and constraints are minimal, with distillation columns having a diameter that will allow them to be transported to the plant and the gasifier being built with durable walls that can withstand high temperatures and pressures for long periods of time. General safety standards will be accounted for and the plant site location will be constrained to locations further from population centers, environmental reserves, wildlife or public water sources. Proximity to water sources such as ponds or lakes will have to be optimized in order to not pollute the water while keeping water transportation costs low. One key safety requirement is that the chimney or release of waste gases has to be high above ground level in order to keep air pollution near the ground low and to disperse the waste gases. 9


In this syngas fermentation process, this requirement is very important as syngas will have small amounts of unwanted or toxic gases such as H2S and methane that will have to be dispersed. Otherwise, these gases will have to be converted to other by products before they are released. IT systems that will be required to maintain safety will include the basic process control systems such as PID controllers as well as MPC controllers. Safety IT systems such as SIS interlock systems will also be put in place in order to lockdown processes such as gasification that can endanger the entire plant if they are out of control. Lastly, the key safety standards for a plant such as containment and emergency procedures will have to be detailed and put in place before the plant begins operating.

2. Existing solutions 2.1 1st generation feedstock One of the leaders in ethanol production from starch along with sugarcane is corn. Currently, about 95% of ethanol in the United States comes from corn due to lower cost and vast research on production (Pimentel & Patzek, 2008). There are two major ways of processing corn into ethanol, namely the dry milling and wet milling process. The dry milling process is more common as it requires less capital to build, is more focused on ethanol production and provides animal feed (dry distillers’ grain) as co-product. On the other hand the wet milling method separates the corn for different uses and is able to produce a variety of product but is also more costly. Over 88% of the ethanol produced in the United States is produced using the dry milling process and the remaining 12% is from the wet milling process (Kwiatkowski, 2006) .An overview of the dry milling process which uses the biochemical process of hydrolysis using enzymes and then fermentation is as follows: ● The corn grain is sent through a series of screen or blowers in order to separate any foreign object such as rocks or minerals. ● The corn then is then crushing and/or grinded and sent to a slurry tank which contains water, enzymes and pH stabilizing chemicals.

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● The mixture is heated and the enzyme break down the starch polymer into shorter chains, in a step called liquefaction. ● The resulting slurry then undergoes hydrolysis which further breaks down the glucose chains into glucose units. The glucose is cooled and undergoes fermentation where it is converted into ethanol with water and carbon dioxide as by products ● The ethanol obtained from fermentation is heated and sent through a degasser drum to flash off the vapour. The resulting products go through a series of distillation, stripping column and molecular sieve in order to separate the ethanol from the rest of the products. The rest of the product separated from ethanol is dehydrated through series of liquid-liquid separation and liquid separation such as centrifuge and dryer (Kwiatkowski, 2006) (Wang, 2007) The capital cost of a corn ethanol production plant with capacity of 400 ML/year will be $220 million (using CEPCI to find value in present value of 2014) per plant and the cost of corn will be $725 million and utility costs of $642 million. Through research, corn ethanol production and use could reduce GHG emissions by 18% of current levels. This however does not account for the deforestation of land in order to grow more corn since it is a crop that requires soil with high nutrient concentration. While corn ethanol is a mature industry, it continues to face issues of minimal greenhouse gas emission reduction, negative net energy balance and decrease in corn food supply. Corn is a big part of human food consumption and the use of corn as fuel often become an ethical issue. As a result, intense research on cellulosic ethanol lead to the discovery of second generation feedstock which is also known as lignocellulosic feedstock.

2.2 2nd generation feedstock Second generation feedstock takes advantage of the abundance of biomass on the planet. Second generation feedstock uses cellulose and hemicellulose which are complex

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sugar polymers found in natural biomass such as grass, wood and trees as source of ethanol. In terms of feedstock, many types of biomass can be used since all of them contain cellulose. Switchgrass and Miscanthus are mostly studied and used as the feedstock. Switch grass is a perennial grass native to North America and its abundance make it appealing to be used as feedstock for ethanol production.. A few examples of second generation feedstock include Switchgrass and Miscanthus where Miscanthus was the preferred feedstock in the bioethanol plant design. Unlike corn the use of switchgrass or miscanthus as feedstock has no impact on food supply and is therefore more appealing. Furthermore, growth of miscanthus and/or switchgrass requires lands with little to no fertilizer, pesticides or energy input which is opposite to that of corn. Preference of using miscanthus as the bioethanol plant feedstock over corn and other second generation can be seen by looking at greenhouse gas emissions associated to land conversion for increased corn, switchgrass and miscanthus growth. A study by Mueller et. al has shown that CO2 equivalent emissions from corn ethanol plant in the U.S. is rated at 92 g CO 2 equivalent per MJ energy provided which is a marginal benefit to gasoline’s 96 g CO2 equivalent per MJ energy provided (Dunn, 2013). The same study also indicate that greenhouse gas emission from land conversion for increased production of corn ethanol is highest at 7.6 g CO2e/MJ while Miscanthus has the lowest at -10 g CO2 equivalent per MJ due to its carbon sequestrating ability and high yield (Dunn, 2013). Furthermore,. Dunn et.al has shown that the average peak annual biomass of miscanthus is 22 tonnes of biomass per hectare while switch grass only produced 10tonnes of biomass per hectare. The same study also shows that the yield of miscanthus is less sensitive to the amount of rainfall and fertilizer compared to switch grass (Dunn, 2013). Field trials in three locations the United States have shown that miscanthus yield is three to four times of that of switchgrass (Liska, 2009). A side by side comparison of switchgrass and miscanthus greenhouse gas emission reveal that emissions produced by Miscanthus growth, harvesting and transportation is about 31% lower than that of switchgrass.

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In second generation ethanol production, lignocellulosic ethanol has many advantages over first generation including lower GHG levels and abundant feedstock supply. However, this technology is not commercialized yet due to high capital and operating costs on some of the process components as well as the enzymes/bacteria used can be expensive. The steps to lignocellulosic ethanol production include pretreatment, hydrolysis and fermentation. This can be seen in Figure 1.

Figure 1. A hydrolysis-based cellulosic ethanol production process. (Dwivedi et al., 2009)

In the pre-treatment step, the lignin walls of the biomass is broken down or pushed apart in order to expose the cellulose in order to undergo hydrolysis and fermentation. This step requires high amounts of energy due to the strength in the walls and is the most expensive step and the hardest step in the whole process. Pre-treatment can be done physically, chemically and biologically with the chemical method currently being the most common. One way of physical pre-treatment is done using liquid hot water where high temperature and pressure water is used to breakdown the lignin walls. This method has also shown improvements in the sugar recovery as well as partial hydroxylation of the cellulose in the biomass. Another chemical treatment method is

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Ammonia Fiber/Freeze Explosion (AFEX). The biomass is exposed to liquid ammonia at high temperature and pressure, and then a swift reduction in pressure exposes the cellulose which can then be processed. Other types of chemical treatment include alkali, ionic liquid and dilute acid treatments. Out of the three types of treatment, biological methods are much preferred due to their ability to produce higher yields (theoretically) while having faster breakdown times and lower emissions. However, biological methods are some of the most expensive and furthest away from commercializing methods out of the three types of pretreatment. Biological pretreatment uses enzymes to breakdown the lignin cell wall to expose the cellulose (Alvira, 2010). After pretreatment, the cellulose undergoes hydrolysis through a biochemical, thermochemical or a combination of both biochemical and thermochemical process which will be referred to as the hybrid process. All three methods will be explained in the following section of this report. The estimated capital cost for a cellulosic plant of 400 ML/year was found to be around $496 million (using 2006 prices and scaled to 2014 present value using CECPI and assuming linear relationship between cost and production capacity) and operating cost of $249 million/year which includes raw materials such as feedstock and enzymes ($102 million), utilities such as water, electricity and maintenance ($54 million) as well as other charges. this estimation is lower than the costs for a corn ethanol plant. It was also found through research miscanthus is able to reduce GHG emissions by up to 88% of current biofuel production.

3. Design alternatives Miscanthus belongs to the second generation feedstock of ethanol known as lignocellulosic ethanol. Miscanthus can be converted into ethanol through the biochemical, thermochemical or the hybrid process which is a combination of both biochemical and thermochemical process. The hybrid process is the recommend process for the bioethanol plant design due to several advantages over the biochemical process and thermochemical process which are highlighted below. A brief description of each

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process is presented and a comparison between each process is made. Finally, proposed location of the bioethanol plant is also presented at the end of the section

3.1 Hybrid process The hybrid process combines the thermochemical pathway of gasification of feed stock into syngas with the biochemical pathway of fermenting syngas into ethanol. A general step of producing ethanol from Miscanthus using the hybrid process is as follows;

1. Drying and crushing 2. Gasification 3. Gas cleaning and cooling 3. Fermentation 4. Distillation/purification

In the pre-treatment step, raw feedstock is dried to a moisture content of 10% water. The dried feedstock is crushed and introduced to a gasification reactor where steam is also introduced. The heat from steam disintegrates the feedstock into its elemental components. A series of exothermic reactions occur and heat the gasification reactor to around 850 C (Dwivedi, 2009). The reactions are also responsible for the production of carbon dioxide, carbon monoxide, hydrogen and trace amounts of hydrogen sulfide, ammonia and methane - a mixture gases known as syngas. Other products from the gasification step include solids such as ash and char. Syngas undergo a series of gas cleaning and gas cooling steps where any impurities like hydrogen sulfide, ammonia and methane are removed and syngas is cooled to 37°C for fermentation.. Equipment used to remove impurities in syngas may include, adsorption column, scrubbers or cyclone for solids removal The cleaned syngas is then sent to a fermentation vessel where bacteria such as Clostridium Ljungdahlii anaerobically digests syngas into acetic acid, ethanol and water at 37 C and 1 atm (Abubackar, 2011). The product from fermentation is a combination of ethanol, acetic acid and water. The fermenter product is then sent to a

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series of separation sections such as distillation column and molecular sieves in order to purify the ethanol.

3.2 Biochemical process Production of ethanol from Miscanthus employs the biochemical process of hydrolysis using enzymes and then fermentation similar along with several pre-treatment steps. A general step for producing ethanol from Miscanthus is as follows: 1. Pre-treatment 2. Enzymatic hydrolysis 3. Fermentation 4. Distillation/purification First the raw feed stock of miscanthus or any second generation undergo drying where moisture content is generally brought down to approximately 10%. The dried feedstock is then crushed to a size of approximately 3.2 mm. The crushed feedstock goes through a series of pre-treatment steps as outlined in section 2. Once the cellulose and hemicellulose are rid of lignin and can be exposed to enzymes, enzymatic hydrolysis proceeds. Enzymatic hydrolysis is the process where the long polymer sugar chains which makes up of cellulose are broken down into sugar monomers such as glucose, fructose and xylose. The resulting monomers are then metabolically digested by bacteria under anaerobic conditions where alcohols such as ethanol are produced. The products which consist of several long chained alcohols, acetic acid and water are then sent to a series of separation steps in order to purify the ethanol. The biggest difference between the biochemical and hybrid process is the need of a pre-treatment step in the biochemical process. Pre-treatment is energy intensive and is a huge drawback to the biochemical process of ethanol production. In contrast, the hybrid process does not require pre-treatment and is therefore less energy intensive which results to lower operating costs. It is projected that about 20% of total cost of cellulosic ethanol

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production is from pre-treatment, a cost that is non-existent in the thermal-biochemical process (Khanna, 2008). The type of pre-treatment employed is also dependent on the type of feedstock for cellulosic ethanol. For example, using corn stover as feedstock uses a different, less expensive pre-treatment process as compared to using switchgrass, which benefits from ammonia fire explosion pre-treatment. This results to lesser feedstock flexibility for an ethanol plant using the biochemical process and played an important in the group’s decision of using the hybrid process. Another important factor issue would be fully breaking down the grass, as the pre-treatment stage is not as effective on grasses with high lignin contents like Miscanthus (~23% lignin content) (Sanchez , 2008). Studies also reveal that the choice of pre-treatment has an effect on upstream processes (i.e. harvesting and storage) since aging of the feed stock during storage can make it resistant to certain types of pre-treatments. Furthermore, the choice of pretreatment has great effects on the downstream processing. In the thermal-biochemical hybrid process pre-treatment is completely eliminated and as a result upstream process such as harvesting and storage has very little effect on downstream processes. Another advantage of thermal-biochemical process over the biochemical process is the increased ethanol yield associated with the thermal-biochemical hybrid process. In the biochemical process lignin is often unused and separated in the pre-treatment process. In contrast, the hybrid process utilizes the whole biomass including the lignin in the gasification process. Furthermore, a significant portion of 5-carbon sugars from hemicellulose cannot be completely converted into alcohol and better enzyme technology is needed (Daniell, 2012). This results to lower ethanol yield per tonne of feedstock using the biochemical process. Finally, pre-treatment in biochemical process is a relatively new technology and research is currently ongoing. On the other hand, gasification in the hybrid process is a much older technology and is used in processes besides ethanol production. In addition to the pretreatment step, the biochemical process requires enzymatic hydrolysis which breaks down the network of polymers that make up cellulose and hemicellulose into sugar monomers. The two main types of hydrolysis are either acid or enzymatic hydrolysis. The downfall of acid hydrolysis is it produces inhibiting

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microorganism which lower ethanol yield in the fermentation stage. Acid hydrolysis also causes corrosion of equipment and the acid needs to be recovered at the end of the process. In the hybrid process inhibitors are not present since acid hydrolysis does not occur which allows for a more consistent ethanol yield. Furthermore, acid is not involved in the hybrid process and therefore lower corrosive material can be used. In enzymatic hydrolysis, enzymes such as cellulose break down cellulose and hemicellulose into sugar monomer units. A drawback of enzyme hydrolysis is the cost and need for large scale production of enzyme. A table of cost of a cellulosic ethanol plant producing 58 M gallons/year of ethanol using the biochemical process is shown in Table 1. Table 1. Economic analysis of a cellulosic ethanol plant producing 58 M gallons/year of ethanol using the biochemical process

Process Section

Cost (Millions $U.S. 2013)

Feedstock handling [1]

14.5

Pretreatment[2]

47.9

Xylose fermentation

12.5

Enzyme production [1]

5.7

Saccharification and fermentation

42.2

Ethanol recovery[3]

8.1

Utilities

102.6

Total

233.5

Table reproduced from Foust,2009 and inflated to 2013 dollars using CEPCI index.

3.3 Thermochemical process Besides the biochemical and hybrid process, lignocellulosic ethanol is also produced through thermochemical process which converts syngas produced in the

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gasification step into ethanol and other linear alcohols using a synthetic catalyst. A general step of producing ethanol from the thermochemical process is as follows:

1. Drying 2. Gasification 3. Syngas cleaning 4. Catalytic conversion of syngas into ethanol and alcohol 5. Distillation/Purification

First, the feedstock is removed of impurities through washing. The washed feed stock is then dried and grinded/crushed into smaller pieces. The feed stock is then fed to a fluidized bed gasifier and can reach high temperatures (800°C). Due to high temperature, the feedstock decomposes to syngas which is made up of carbon monoxide, carbon dioxide and hydrogen. The syngas is collected from the top of the gasifier and is cooled through a series of heat exchangers. The cooled gas undergoes water scrubbing steps where tar and residuals are removed. The gas is compressed to a higher pressure and impurities such as hydrogen sulphide and carbon dioxide is removed in an amine unit. The cleaned gas is sent through a bed of fixed bed molybdenum disulphide based catalyst which produces ethanol along with other linear alcohols. The mixture is sent through a series of distillation and separation steps where the ethanol is obtained (Yang, 2008). Compared to the hybrid process, it is evident that the pure thermochemical process result to several by products such as methanol and other linear alcohols which require several separation steps. In the hybrid process, the main products are ethanol, acetone and water (bacteria media) which requires fewer separation units. Another advantage of the hybrid process is that conversion of syngas to ethanol occurs at low pressures (1 bar) and low temperatures (37°C) which results to lower operating costs. A study has also shown that the bacteria used in the hybrid process is also able to tolerate sulfur impurities in the syngas which results to lower energy and cost allocated in the gas cleaning step (Roy, 2014).

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In the thermochemical process, sulfur impurities must be eliminated before the catalytic conversion to ethanol since sulfur irreversibly poisons the catalyst. This also serves to be potentially cost saving since replacing a poisoned catalyst is not an issue in the hybrid process. Finally, ethanol yield in the thermochemical process is very sensitive to hydrogen to carbon dioxide ratio. In order to achieve optimum hydrogen to carbon dioxide ratio of the syngas, a water-gas shift reaction step is normally employed which requires the use of another reactor and more steam input (AdvancedBiofuelsUSA, 2011). In the hybrid process, the hydrogen to carbon dioxide is not needed since hydrogen to carbon dioxide is less of an issue. This results in lower operating and capital costs for the hybrid process. Shown in table 2 is an economic analysis of a cellulosic ethanol plant producing 58 M gallons/year of ethanol using the thermochemical process. Table 2. Economic analysis of a cellulosic ethanol plant producing 58 M gallons/year of ethanol using the thermochemical process

Process Section

Cost (in Millions $U.S. 2013)

Feedstock handling

32.1

Catalyst

2.8

Gasification

34.9

Gas cleaning

84.7

Separation

9.28

Utilities

67.5

Total

231.3

Table reproduced from Daniell, 2012 and inflated to 2013 dollars using CEPCI index.

3.4 Bioethanol Location Due to Ontario’s geographic location, all of Ontario experience climate that is well below the freezing point. While this feature is generally unattractive for crop growth, some regions of Ontario do enjoy warmer than others. Ontario can be split into 5 different 20


regions which are 1)Eastern (Ottawa), 2)Central (Hamilton/Toronto), 3)Southwestern (Sarnia), 4)Northeastern (Timmins) and 5)Northwestern (Thunder Bay) . Northeastern and Northwestern Ontario experience longer and colder climates with annual average temperature of around 8C. On the other hand, Southwestern, Eastern and Central Ontario enjoy warmer climates with average annual temperatures of 12⁰C. This leaves 3 possible regions of Southwestern, Eastern and Central Ontario (Hamilton/Toronto) as possible location for the proposed bioethanol plant. Based on Figure 2 we can see that as we move towards Southern Ontario, the length of growing season increase. This implies that the Sarnia, Windsor and Hamilton/Toronto area is a more preferred region than Eastern Ontario where length of growing days of less than 170 days can be observed. It is expected that the Hamilton and Toronto region is generally unfavourable to the approval of a bioethanol plant due to highly dense residential area. Finally, several bio refineries such as Suncor refinery already exists in the Sarnia region which makes it the preferred location

Figure 2.Length of Growing Season in Ontario.( Agriculture and Agri-Food Canada, 2014).

21


4. Overview of Proposed Process Design 4.1 Process Summary The proposed process for producing Ethanol is referred to as Syngas Fermentation and is considered a hybrid approach for converting biomaterials into fuel. The term hybrid is used because the process incorporates thermochemical aspects such as gasification with biochemical ones like fermentation. The process uses five main stages to turn any lignocellulosic biomass (i.e. switchgrass, miscanthus, wood chips) or biowaste (i.e. corn stover) material into ethanol to be used as fuel.

A block flow diagram

shown below in Figure 3 outlines the process, and a full process description follows.

Figure 3. Block flow diagram of Syngas Fermentation process

Fermentation has the advantage of operating at low temperatures (~37ºC) and pressures (~1 bar) compared to alternative processes, lowering overall energy costs for production. Syngas fermentation also has the advantage of high selectivity of ethanol (leading to increased yields) and good tolerance to typical syngas impurities such as sulfur, which in turn reduces costs for syngas cleaning (Daniell, 2012). Additionally, syngas fermentation operability is not impacted greatly by the H2:CO ratio of the syngas, meaning that the gasification process and the proceeding syngas cleaning steps are awarded flexibility (Daniell, 2012). Lastly syngas fermentation allows for a very large variety of feedstock to be used to produce ethanol, capable of converting virtually any lignocellulosic material into ethanol. 22


4.1.1 Pre-Treatment The raw or harvested feedstock is first pretreated through washing, drying and

crushing stages to bring the biomass into a desirable pellet form, ideally with diameters in the range of 3-6 mm (Roy, 2014) (Michel, 2011). The raw biomass will likely be stored on site in bales (if grassy biomass) or in large storage containers (if woody biomass) before being dumped/placed onto a conveyer belt which initiates the pre-treatment process that follows. The washing stage is a precautionary cleaning stage meant to clean the feedstock of any dirt or lingering chemicals such as pesticides. This pre-treatment stage can be done through a variety of methods, but a conveyer-belt spraying unit was selected for this process. The next pre-treatment step is drying, which is achieved using a belt drying unit, operated using excess steam or air as the drying force (Li, 2012). The belt dryer is simply a conveyer belt pushing the biomass through a unit that continuously dries the materials as they pass.

Drying is a necessary pre-treatment step because the moisture within the

biomass takes away energy from the gasifier which cannot be recovered at an approximate rate of 2260 kJ lost per kilogram of moisture (Basu, 2013).

The total

moisture content of the biomass should be between 10-20% ideally for minimal energy loss (Roy, 2014) (Basu, 2013). Next, the dried biomass must go through a size reduction step for ease of loading, and for optimal performance within the gasifier. A continuously operated hammer mill crushing device is used in this process, selected for its ease of operation and control of desired particle size while being able to handle a variety of different feeds (Kratky, 2010). For optimal operating conditions the biomass moisture content must not exceed 10-15% (Kratky, 2010). The crushed dried pellets that are left from the pre-treatment process are sent towards the gasifier via conveyer belt.

23


4.1.2 Gasification Once pretreated, the biomass is fed into a gasifier unit where it is converted into a

syngas mixture, composed mainly of carbon monoxide (CO), hydrogen (H2), carbon dioxide (CO2), methane (CH4), and impurities.

Many types of gasifier configurations

may be employed for this task, such as counter/co-current fixed bed, fluidized bed, or entrained flow gasifiers. For the purpose of biomass gasification a fluidized bed gasifier was selected, mainly due to its insensitivity to fuel quality, allowing for flexibility in the biomass feedstock (Basu, 2013). The fluidized bed gasifier is identified through its use of a “bed”, which is essentially a collection of granular solids that are kept suspended via the continuous flow of gases at specific velocities (Basu, 2013). The bed, selected as quartz sand, provides excellent solid-gas mixing and a relatively uniform temperature profile within the gasifier (Basu, 2013). More specifically, the fluidized bed gasifier is a circulating fluidized bed gasifier, where the bed is recirculated within the gasifier providing longer gas residence times and allowing for larger units in general (Basu, 2013). An image of a typical circulating fluidized bed gasifier can be seen below in Figure 4. The gasifier is operated at low pressures (~1 bar) and high temperatures (~800-1000C) and is naturally an energy intensive process. Most of the lost energy in operating the gasifier can be recovered downstream as heat through cooling of the syngas.

24


Figure 4. Typical fluidized bed gasifier configuration (Basu, 2013)

Within the gasifier, the biomass feed undergoes incomplete combustion, in the presence of either air or trace amounts of oxygen, to produce syngas mainly through the following reactions.

 + →  +→ + +→ +

[1] [2] [3]

Reactions 1-3. Reactions taking place within g asifier. Incomplete combustion is achieved with a controlled amount of oxygen.

The carbon containing biomass is fed into the gasifier with steam and a controlled amount of oxygen that ensures the biomass undergoes incomplete combustion, starting a chain of reactions resulting in the final syngas mixture. The hot syngas is continuously

⁰

drawn from the gasifier at an approximate temperature of 850 C, as is the produced ash from the bottom of the unit.

25


For the purposes of feeding the biomass to the gasifier there are a variety of different methods, but a gravity chute was selected for this process due to its simplicity. After pre-treatment the feed is dropped onto a feed screw which leads to a gravity chute that feeds the biomass directly into the gasifier (Basu, 2013). A schematic of this feed system is shown below in Figure 5.

Figure 5. Typical biomass feeding system for fluidized bed gasifier .(Basu,

2013)

The feed screw allows for relatively simple control of feed flow, and the gravity chute offers a simple method of feeding biomass to the gasifier. The tip of the gravity chute lies within the gasifier itself and must be properly insulated to withstand the high temperatures within the unit. For this feed configuration the unit is often operated at slightly below atmospheric pressure to ensure that the rising gas doesn’t travel into the feed chute (Basu, 2013). A jet vapour stream placed directly under the chute is often installed to ensure that no gas travels up the chute.

4.1.3 Gas Cleaning The syngas produced from the gasifier is a gaseous mixture containing H2, CO,

CO2, CH4 and a multitude of impurities. Based on further downstream processes, the

26


ideal syngas composition should be low in impurities and high in H2/CO content for the purposes of fermentation. Certain impurities such as hydrogen sulfide (H2S), ammonium (NH3), and hydrochloric acid (HCl) can cause issues downstream if their levels are too high. Syngas cleaning is categorized into two types, hot-gas cleaning (HGC) or cold-gas cleaning (CGC). Attached to the gasifier is a cyclone which quickly removes any solid particulates or ash within the syngas before it undergoes further stripping. A cyclone is a simple way to screen out solid impurities and can be operated at temperatures up to

⁰

~1000 C (Basu, 2013). The hot syngas mixture leaving the cyclone is next sent into an adsorption column

used primarily for removing H2S which can cause potential issues during downstream processes and sulfur has been known to corrode metal surfaces (Woolcock, 2013). First

⁰

the syngas must be cooled to an approximate temperature of 600 C for the adsorbent within the column, zinc oxide in this case, to work effectively (Woolcock, 2013). During the cooling of the syngas some of the heat lost in the gasification process may be recovered as steam, which can be fed to a steam turbine to produce electricity. The adsorption column is packed with an iron oxide adsorbent which selectively binds with the sulfur particles to form a metal sulfur compound such as ZnS or FeS (Woolcock, 2013). For this process a Zinc oxide (ZnO) adsorbent was selected mainly due to its low cost and high availability. The reversible adsorption columns are to be run in parallel, with one column running at a time.

When the sulfur compounds fully bind to the

adsorbent bed the flow is sent to the parallel adsorption column.

The fully bound

adsorption column is then fed a stream of oxygen which regenerates the bed by unbinding the sulfur back into the gaseous stream (Woolcock, 2013). A gas rich in sulfur dioxide exits the regenerated adsorption column where it is sent to a sulfur recovery unit to obtain elemental sulfur or sulfuric acid (Woolcock, 2013). The parallel configuration of the adsorption columns ensures that the process may be run continuously as one bed is being regenerated while the other one is in operation. Following the adsorption column the syngas undergoes rigorous cooling stages to

⁰

reach an approximate temperature of 45 C, all the while recovering significant amounts

27


of energy as heat. Having now transitioned into cold-gas cleaning, a wet scrubber is selected to remove the remaining particulates. Ammonium and chlorine are both highly soluble in water making the wet scrubber a great choice in removing these impurities (Woolcock, 2013). The gas enters a column that is known as a spray tower, which is essentially a vessel that contains porous pipes that spray the passing gas with water which collects the impurities. Along with removing ammonium and chlorine the wet scrubber will also remove any leftover or newly formed solid particulates in the gas. The water is continuously drained from the bottom the tower and sent to wastewater treatment for processing. A typical configuration of a wet scrubber is shown below in Figure 6.

Figure6. Typical wet scrubber configuration (Woolcock, 2013).

⁰

After leaving the wet scrubber the syngas temperature has fallen to approximately

37 C which is the ideal temperature required for the fermentation step that follows (Roy, 2014). The gaseous mixture is also free of the impurities that could cause problems downstream and is ready to be converted into ethanol via microbial fermentation.

28


4.1.4 Fermentation The clean syngas is now free of impurities and cooled to a low temperature perfect

for the fermentation process that follows. The syngas is fed to the fermenter where it aids in the production of ethanol. Essentially the syngas is converted to ethanol via a number of reaction pathways that are made possible by certain strains of bacteria. To date, the most relevant family of bacteria utilized in syngas fermentation is the Clostridium family (Abubackar, 2011) (Daniell, 2012). Of the various strains within the family, Clostridium Ljungdahlii is the most widely studied and is used in this process primarily for its ethanol selectivity properties. Isolated primarily from chicken farm waste, the main challenge of the process would be obtaining the bacteria, as it is not easily isolated (Abubackar, 2011). Though there are pilot-scale and pre-commercial plants in operation that use these bacteria, demonstrating the feasibility of the process (Daniell, 2012). The role of the bacteria in the process is that of a biocatalyst, it enables certain reactions to occur, while the bacteria itself is hardly consumed (Abubackar, 2011). The bacteria can then be regenerated or recycled to maximize process efficiency and minimize bacteria losses. For optimal growth of the bacteria the temperature of the reactor should

⁰

be held as close to 37 C as possible with the pH maintained at 6, though acceptable performance can be achieved within a pH range of 4-7 (Roy, 2014) (Abubackar, 2011) (Daniell, 2012). The bacterium has also been shown to support growth on ethanol, further improving the overall bacteria efficiency (Daniell 2012). This bacterium and others of the clostridium family enable the syngas to take the Wood-Ljungdahl biochemical pathway in order to produce ethanol. shown below in Figure 7.

29

A simplified reaction pathway is


Figure 7. Wood-Ljungdahl biochemical pathway (Abubackar, 2011).

In short, the CO and H2 are utilized as the main reactants for ethanol production and their amounts in the syngas dictate which reactions are utilized (Daniell, 2012). The main reactions that take place within the fermenter, made possible through the biocatalyst are shown in Equations 4-10, forming ethanol (CH3CH2OH) and acetic acid (CH3COOH) as the main products (Daniell, 2012).

6+3→+4 4+2→ +2 3+3 →+ 2+2 → 2+4 →+ 2 +6 →+3 2 +4 →+2

[4] [5] [6] [7] [8] [9] [10]

Reactions 4-10.Primary reactions that occur within fermenter, dependant on H2/CO content(Daniell, 2012)

30


Initially acetic acid will be the favoured product but once the fermenter is run with recycle for several hours ethanol production will be favoured, reaching an approximate, steady ethanol: alcohol ratio of 2 (Abubackar, 2011). It should be noted that since the bacteria is anaerobic the reactor must be kept free of any oxygen or the bacteria will die (Abubackar, 2011). As with gasification, many fermenter types may be incorporated for the purposes of syngas fermentation and the process can be run in either batch, semi-continuous, or continuous modes of operation dependent on which fermenter is used. The most studied and widely employed reactor configuration for the purposes of syngas fermentation is a stirred-tank bioreactor (STB) and was selected for this process under continuous operation (Abubackar, 2011). A typical configuration of an STB is shown below in Figure 8 .

Figure 8. Typical stirred-tank bioreactor configuration. l- gas sparger; i- gas feed; ii- medium feed; iii pump; iv- liquid outlet; v- gaseous outlet (Abubackar, 2011).

31


The gaseous feed enters the fermenter at the bottom of the reactor where most of the syngas breaks into smaller bubbles, well dispersed by the continuous mixing of the tank. Syngas retention times vary but can be expected to be approximately 1 minute (Abubackar, 2011). The liquid broth from the reactor is continuously drawn as fresh medium is being pumped into the bioreactor. Ideally the syngas will have a carbon conversion efficiency that can reach up to 80%, where the unconverted gas is also removed from the bioreactor continuously (Daniell, 2012). The unconverted syngas, now high in CO2 content, can be combusted to recover even more of the energy that was used up in the gasification step. The fermenter also contains a fermentation medium, which varies greatly in composition, but is largely made up of acidic water. The medium also includes the bacteria, nutrients, vitamins, minerals, salts, yeast extracts and/or other additives that are required for ethanol production. The liquid medium extracted from the fermenter is usually immediately filtered to remove the bacteria, which is recycled, as it would die during the distillation stage. Some of the ongoing challenges with this process are limited mass transfer rates, which can be improved through modification of the bacteria or incorporating a 2-stage system, and limited ethanol concentrations in the fermentation broth (Abubackar, 2011). Typically, the broth can’t contain much more than 5% ethanol before it impacts the bacteria and causes problems within the unit, so the resulting ethanol yield is ~3-6% of the concentration within the broth (Abubackar, 2011) (Daniell, 2012). Ideally the yield from the bioreactor is 0.3-0.4 L ethanol/kg-dry feedstock (Roy, 2014).

4.1.5 Distillation The final broth that is pulled from the fermenter contains mainly the fermentation

medium, with ethanol (~3-6%) and acetic acid (~3-6%) in low concentrations. This broth is then sent to a series of distillation columns followed by extractive dewatering to reach final product purity.

The ethanol is to be separated from the acetic acid and the

fermentation medium using successive distillation columns in series. The bottoms from the first distillation columns can be recycled back to the bioreactor since it will mainly

32


contain the fermentation medium, though a purge stream is necessary to prevent accumulation of acetic acid. All of the energy recovered from cooling the syngas and combusting the unused syngas can now be used to provide energy for the operation of the distillation columns. Upon successive separation of ethanol, oftentimes the ethanol is sent to further separation processes such as dewatering to ensure high product purity to meet quality specifications.

4.2 Design Basis The design basis that was selected to base the sizing, costing and economic analysis for this process was to produce 100 million liters of ethanol per year. Meeting this production rate must be done while maintaining high product purity (>99%) and ensuring safe process operation.

4.3 Product Specifications The final ethanol blend must meet the standards set by the Canadian General Standards Board and the American Society for Testing Materials (ASTM) to ensure safe operation within a motorized vehicle, and other standards outlined further in the report.

5. Process Behaviour 5.1 Normal operation The feedstock (miscanthus) which contains 12-20% moisture content is fed to a continuous belt dryer drier at atmospheric pressure and room temperature where the dried it is dried to a 10% moisture content. Medium pressure steam which enters at a temperature and pressure of 162°C and 7.8 atm respectively and exits at 120°C and 2 atm is used as the heating media. The dried feedstock is then sent into a hammer mill which reduces the feedstock into a target particle size of 3.2 mm. Once the feedstock are milled into the appropriate particle size it is sent into the gasifier unit which operates at 850°C

33


and 1 bar. Steam is also introduced into the gasifier and provides heat to disintegrate the feedstock into its elements. The reactions presented in chapter 4 occur inside the gasifier and the resulting products consist of solids such as char and ash along with the major product that is syngas. Other major by products include ammonia, hydrogen sulphide and hydrogen chloride. The gasifier products are sent to an H2S adsorber column where H2S is removed by adsorbing on to a zinc oxide chemical adsorbent. The zinc oxide adsorbent eventually needs to be regenerated after normal operation, so after about 15 days (based on volume of adsorbent and its capacity), the feed going into the adsorber is sent to a secondary adsorber that operates in the same way. During this time, the used zinc sulfide being regenerated is contacted with oxygen in air to convert it back to zinc oxide through the Reaction 11:

 + 32  →  + 

[11]

Reaction 11. Regeneration of adsorbent bed via oxygen.

The sulfur dioxide produced from this reaction is then sent to a sulfur plant to produce a sulfur compound or for other processing. The sulfur-free syngas is then cooled from its temperature of 550°C to 37°C and is sent into the wet scrubber part of the gas cleaning section, also known as cool gas cleaning. The wet scrubber is simply a vessel where the cool syngas is contacted with water from a water spray to removeCH3, HCl and CO2. Once the impurities are removed, the syngas is continuously sent into the anaerobic (closed-roof) fermentation vessel through an entrance from the top. A recycle stream which contains bacteria and nutrients from a storage vessel is also introduced into the fermentation vessel as a mixed liquid broth at 37°C and 1 bar. At the same time, a continuous feed of fresh bacteria broth which contains nutrients essential for bacterial life is mixed with the recycle broth from the bottom stream of the first distillation column.. The mixed recycle and fresh broth mix to a temperature of 73°C and passes through a shell and tube heat exchanger where the mixture is cooled to 37°C. The pH or acidity of the broth mixture introduced into the fermentation vessel is kept at an optimal pH of 6 by 34


controlling both the amount of bacteria broth recycled into the storage vessel and the amount of fresh bacteria broth. The resulting products from fermentation are ethanol, acetic acid and water which represent bacteria broth at a temperature of 73°C and pressure of 1 bar. The fermenter liquid effluent is then fed into the first distillation column which operates at a condenser pressure of 1 bar and a reflux to distillate ratio of 0.1. The bottom stream of the distillation column has a molar fraction of 95.69% water, 4.30% acetic acid and .0042% ethanol at 100.174°C and 1 atm. The distillation column reboiler uses low pressure steam at inlet conditions of 135°C and 3 atm and outlet conditions of saturated liquid (vapour fraction = 0) and a pressure of 2.7 atm. The bottom stream which is now composed of mostly water and hence bacteria is mixed with fresh bacteria broth. Meanwhile, the distillate stream exits the top of the distillation column at a temperature of 86.95°C and pressure of 1 atm with molar composition of 89.5% water, 1.40% acetic acid and 9.05% ethanol. Cooling water with inlet conditions of 32°C and 1 atm and outlet conditions of 48°C and 1 atm is used to condense the vapour from the top of the distillation column. The distillate stream is then fed into a second distillation column which concentrates the ethanol. The second distillation column operates at a condenser pressure of 1 atm and a reflux to distillate ratio of 0.79 moles. The bottom stream of the second distillation column exits at a temperature of 97.8°C and 1 atm with molar fractions of 97.7% water, 0.73% ethanol and 1.6% acetic acid. The bottoms stream of the second distillation column also uses low pressure steam at inlet conditions of 135°C and 3 atm and outlet conditions of saturated liquid (vapour fraction = 0) and a pressure loss of 0.25 atm. Meanwhile, the distillate stream of the second distillation column exits at a temperature of 79°C and 1 atm with molar fractions of 21.5% water, 78.4% ethanol and less than 0.1% acetic acid. Cooling water which enters at 32°C and 1 atm and exits at 48°C and 0.9 atm is also used to condense the vapour stream of the second distillation column. The distillate stream of the second distillation column is sent to a dehydration process where ethanol with a purity of 99.9% is obtained. Although these conditions are

35


based off of a simulation, the only realistic expected change is a slight increase in the acetic acid concentration in the final ethanol stream leaving the plant.

5.2 Start-up and shutdown Clarity and unambiguity in startup and shutdown procedures is necessary in order for plant operators, engineers and technicians to operate the plant with little variability and a constant throughput. However, to appropriately address this essential part of running a chemical plant, each startup and shutdown should begin with awareness and preparation by operators. In addition, all plant personnel and employees should familiarize themselves with the correct operation and maintenance procedures. There should be a checklist for the startup and shutdown procedures for the whole plant as well as for each individual unit and section of the plant (e.g. for a compressor or cyclone). The operators who will be starting the procedures should be well trained and confident in their ability to handle unexpected circumstances. Inter-personnel communication between plant operators, engineers, technicians and managers is also necessary, and everyone should work cohesively and with specific tasks and objectives in mind, which should be decided upon before the plant starts up or shuts down. Lastly, the working materials such as catalysts, refrigerants or adsorbents and utilities like cooling water or steam should be readily available for the process as needed. The following is only a general operating procedure for the miscanthus syngas fermentation plant, with startup and shutdown sections. 5.2.1 Start-up 1. Run hot pressurized air through the gasifier and it should exit through the flash

column which releases used syngas out of the fermentation tank. This ensures that the syngas exit out of the gasifier is unhindered and all residual gases/materials present in the gasifier and the adsorption column, scrubber and fermentation tank are removed. It also ensures that the valves are working properly and the equipment, pipes and valves are not blocked. Slowly, the miscanthus feed and high pressure steam (used to heat the gasifier for pyrolysis to begin) can be added

36


simultaneously, and can be brought to steady state by changing the throughput from upstream (pretreatment). 2. Startup the hammer mill by first starting up any necessary heaters or burners, bringing these up to the desired operating temperature, and once these are at set point, begin introducing feed slowly, and gradually increase the throughput until the desired steady throughput is achieved. 3. Drying can be started up by running air through the process before bringing in the feed. 4. First, run a test batch of miscanthus through the hammer mill and drying processes to see if the miscanthus water content and size is according to specifications that ensure maximum efficiency through the pyrolysis process in the gasifier. 5. The product from the gasifier exit is not yet run through the gas cleaning processes, so it exits through a flare side stream. 6. For the heat exchangers, open shell side vent valve to release air or gases, slowly introduce cooling fluid until shell side is flooded with cooling fluid, shut the shell side vent valve, open the tube side vent valve to release air or gases, slowly introduce the syngas until all tube (passes) are filled, close the tube side vent valve, slowly increase syngas flow rate up to operating conditions. Any utility processes including condensers, reboilers and heaters should run this process. 7. Fermentation can begin by first running it as a batch process separately before the syngas is produced and sent through. This batch process is required in order to facilitate the conditions necessary for the bacteria to survive, namely a temperature of 37ºC and around 5-6 pH. Once the bacteria has been introduced and the conditions are satisfactory, the process can be made continuous by the introduction of syngas and the release of some broth along with the products (including ethanol) and the used syngas. 8. For the pump (which is used to pump broth and water for fermentation), make sure all connections are in place. Close the discharge valve and open the suction valve. Slowly introduce the water/broth to the pump until the pump suction line

37


fills. Open the discharge valve and start the pump. Once the pump reaches the desired speed, open the discharge valve to a setting that gives the best efficiency point. 9. When the gas cleaning, heat exchanger and fermentation processes are ready, the syngas flow rate to the flare is slowly reduced and instead this flow begins running through the downstream processes. Steady state is achieved when almost no syngas exits through the flare. 10. Operation of the distillation column begins by removing undesirable materials in the column using air or inert gases. The next step is to slowly increase the pressure inside the column using one of the components of the feed or an inert gas. A small amount of the feed is then introduced to the distillation column and the column is ran at total reflux and utilities are turned on. Gradually bring the column into normal operating conditions. 5.2.2 Shutdown 1. The pretreatment processes including the hammer mill and the drying need to be

shut down simultaneously by first reducing the air into the dryer and subsequently lowering the feed rate. 2. Let the remaining pyrolysis reactions happening in the gasifier finish after there is no feedstock entering the gasifier. Gradually reduce the steam flow rate into the gasifier. Shut down the gasifier after all the miscanthus has burned and the char and particles at the bottom of the gasifier are cleaned out. 3. Shut down the cyclone, adsorber and wet scrubber when there is no feed. 4. Shut down the water inlet into the fermentation tank when the last syngas has reacted in the tank. Simultaneously slow down the steam inlet and cooling water inlet into both distillation columns’ reboilers and condensers. Slowly, the feed will exit through the bottoms of the second distillation column. The recycle stream will still operate until all the broth and water has exited from the second distillation column. This entire step ensures that the ethanol product stream does not receive any of the other components.

38


5. Drain the entire system of broth, wastewater and ethanol. 6. Clean all units and removal residual materials.

5.3 Emergency Procedures 1. In the event of an emergency, response teams as well as authorities must be contacted immediately. This will include on site response as well as off site (ie, 911). 2. Evacuate the site of emergency to ensure minimum amount of damage and potential casualties. 3. In the event of temperature emergencies, the equipment would be shut down immediately and the energy cooling system would started in order to contain the heat. a. If the heat issue continue, shut down the whole plant to make sure heating is not coming from another part of the plant. b. If a quench system is available on plant, initiate as soon as equipment is shut down. 4. In high pressure emergency situations, shut down the equipment in question and sections prior in attempt to reduce the pressure. a. Open all of the relief valves close to the high pressure area b. If the pressure continue to increase on larger units pass critical point, alert all parties in and around the plant and evacuate as quickly as possible c. In case of pressure being too low in a part of the plant to the point of causing a vacuum, the same procedures as the high pressure emergency situation can be used. 5. Due to the high temperature nature of the gasifier and the streams/units afterwards, any leaks in that part of the plant may result in large amount of damages. a. Shut down the gasifier to stop the production of high temperature syngas. b. Initiate quench system and/or emergency cooling system if available. c. Evacuate the plant floor 39


d. In case of fire in the plant, activate the fire alarm and alert authorities. d. Shut down the plant and evacuate. e. Allow authorities and emergency response teams to resolve the issue. 6. In case of spilling/leaking from pipes or units (liquid or vapou r), the spill/leak must be located and shut down or the process must be re-routed. a. Evacuate areas of spill. b. Any spills must be cleaned up as soon as possible to prevent any potential chemical damage or fires c. Clean up spills using spill pillows or equivalent substance which can absorb the spill d. In case of gas leak from the process, make sure the plant is fully sealed in order to ensure the gas does not escape to the atmosphere e. Check for toxicity before returning to the plant after spill incident f. If contacted with liquid or vapour substance, use proper treatment methods to disinfect contacted area.

6. Overall Material and Energy Balances 6.1 Overall Material Balance: Table 3 displays the overall material balance of the process streams of the bioethanol plant starting from the pre-treatment section which includes drying and milling, to the gasifying unit, gas cleaning section, fermentation unit and the ethanol separation section. Table 3 includes process stream names and the total mass flows in kg/hr for each stream entering and exiting the bioethanol plant.

40

6.1.1 Process Side Table 3. Process side material inflows and outflows of the bioethanol plant PROCESS SIDE – OVERALL MATERIAL BALANCE INFLOW Component

Raw

Steam

Mat Total

122,728

Total mass flow

93,992

247,760

OUTFLOW

Fresh

O2-

NH3

Broth

H2O

Removal

31,040

70,846

735

Waste

WWT

DC-2

Ethanol

Bottom 127,650

952

38, 406

8,819

Evaporated

Cyclone

H2S

Water

underflow

removed

17

218

247,762

6.1.2 Utilities side

Table 4 and Table 5 displays the overall material balance of the utilities used throughout the bioethanol plant starting from the pre-treatment section, to the gasifying unit, gas cleaning section, fermentation unit and ethanol separation section. Table 4 includes the utility material inflows while Table 5 includes the utility material outflows. All numbers have units of kg/hr. Utilities used throughout the bioethanol plant include medium pressure steam (DRYER-STEAM) utilized in the dryer with inlet conditions of 163°C and 7.8 atm and outlet conditions of 120°C and 2 atm. Meanwhile, cooling water (CW1, CW2, CW3 ,CW4) are used for all heat exchangers with inlet conditions of 37C and 1 atm. Low pressure steam (stream name of BOTTOMS1 and BOTTOMS2) with inlet conditions of 135°C and 3 atm is used for the reboiler of the distillation columns while cooling water (DIST1-CW and DIST2-CW) at inlet conditions of 37°C and 1 atm is used for the total condenser of the distillation columns. Cooling water (CW1 and CW2) exit heat exchangers E-101 and E-102 as low pressure steam (stream

119


name of LPS1 and LPS2 respectively) at conditions of 135°C and 3 atm. On the other hand, streams CW3 and CW4 exit heat exchangers E-103 and fermenter cooling jacket respectively at conditions of 49°C and 1 atm. Low pressure stream (BOTTOMS1-LPS and BOTTOMS2-LPS) exit as saturated liquid (vapour fraction of zero) and a pressure drop of 0.3 atm was assumed. Table 4. Utilities side overall material inflows of the bioethanol plant UTILITIES SIDE – OVERALL MATERIAL INFLOWS

Component

Total

DRYERSTEAM 69,945

CW1

CW2

CW3

153,766

89,667

205,641

DIST1-CW

CW4

6,469,180

Total mass

1,490,110

DIST2CW

827,897

BOTTOMS1LPS 60,645

BOTTOMS2-LPS

27,262

9,520,617

flow

Table 5. Utilities side overall material outflows of the bioethanol plant UTILITIES SIDE- OVERALL MATERIAL OUTFLOWS

DRYERComponent

STEAM

LPS1

LPS2

CWR3

CWR4

EXIT Total

69,945

153,766

89,667

205,641

6,469,180

DIST1-CW

DIST2-CW

RETURN

RETURN

1,490,110

9,520,617

Total mass flow

42

827,897

BOTTOMS1 -LPS RETURN 60,645

BOTTOMS2-LPS RETURN

27,262


6.2 Overall Energy Balance Table 6 displays the overall energy balance of the process streams of the bioethanol plant starting from the pretreatment section which includes drying and milling, to the gasifying unit, gas cleaning section, fermentation unit and the ethanol separation section. Table 6 includes process stream names and the total enthalpy flow in kg/hr for each stream entering and exiting the bioethanol plant. 6.2.1 Process Side Table 6. Overall energy balance of the process streams PROCESS SIDE – OVERALL ENERGY BALANCE INFLOW Energy

Raw Mat

Total Fluid

-356,727

Steam

-1,099,100

OUTFLOW

Fresh Broth -495,460

O2-H2O

-16,262

NH3

Waste

Removal -1,963

-61,705

WWT

-15,014

DC-2 Bottom -575,230

Ethanol

-52,071

Evaporated

Cyclone

H2S

Water

underflow

removed

-10,382

-217,910

Enthalpy Flow

6.2.2 Utility Side Table 7and Table 8 displays the overall energy balance of the utility streams of the bioethanol plant starting from the

pre-treatment section which includes drying and milling, to the gasifying unit, gas cleaning sectio n, fermentation unit and the

43

-1.7


ethanol separation section. Table 8 includes the incoming utility stream names while Table 8 includes the outgoing utility stream names and the total enthalpy flow in MJ/hr for each stream entering and exiting the bioethanol plant.

Table 7. Overall energy inflow of the utility streams . All values are in MJ/hr, ex cept temperature which is in units of °C and pressure in units of atm UTILITIES SIDE: ENERGY INFLOW Energy

Total Fluid Enthalpy Flow

CW1

CW2

CW3

CW4

DRYER-STEAM

-2,451,400

-1,427,900

-3,278,400

-103,130,000

73,485

Energy

DIST1-CW

DIST2-CW

DIST1-LPS

DIST2-LPS

Fluid Enthalpy Flow

-23,737,000

-13,188,000

-937,570

-421,480

Electricity

Fluid Enthalpy Flow

ELEC

.0025

Electricity

Energy

MILLING-

P-101

P-102

ELECTRICITY

ELECTRICITY

8x10-4

1.3x10-3

P-103

P-104

P-105

P-106

P-107

ELECTRICITY

ELECTRICITY

ELECTRICITY

ELECTRICITY

ELECTRICITY

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

44


Table 8. Overall energy inflow of the utility streams . All values are in MJ/hr, except temperature which is in units of °C and pressure in units of atm UTILITIES SIDE: ENERGY OUTFLOW DRYEREnergy

LPS1

LPS2

CWR3

CWR4

STEAM EXIT

Fluid Enthalpy Flow

-2,356,900

-1,384,700

-3,260,300

-102,570,000

Energy

DIST1-CW RETURN

DIST2-CW RETURN

DIST1-LPS RETURN

DIST2- LPS RETURN

Fluid Enthalpy Flow

-23,633,731

-13,130,424

-1,069,866

-480,871

69,302

6.3 Stream and Equipment tables 6.3.1 Process Side

Table 9 and Table 10 show the stream tables and equipment table respectively for the pre-treatment section of the bioethanol plant. All mass flows are per mass basis with units of kg/hr while total enthalpy flows are in units of MJ/hr. Pre-treatment section stream table

45


Table 9. Stream table of process streams entering and exiting the pre-treatment section of the bioethanol plant .

Raw Mat.

Evaporated Water

Dried Feed

Miller Feed

Milled Feed

C

55,689

0

55,689

55,689

55,689

H2

523

0

523

523

523

N2

604

0

604

604

604

O2

64,776

0

64,776

64,776

64,776

H2O

28

17

11

11

11

CL2

789

0

789

789

789

S

101

0

101

101

101

Solids

218

0

218

218

218

122,728

17

122,711

122,711

122,711

-356,727

-10,382

-350,498

-350,498

-350,498

Component

Total Flow Total Flow

Enthalpy

46


Table 10. Material and energy inflow and outflow to the equipment of the pre-treatment section of the bioethanol plant

Component

PT-101 [BELT DRYER] IN

OUT

PT-102 [HAMMER MILL] IN

OUT

C

55689

0

55689

55689

55689

H2

523

0

523

523

523

N2

604

0

604

604

604

O2

64776

0

64776

64776

64776

H2O

28

17

11

11

11

CL2

789

0

789

789

789

S

101

0

101

101

101

Solids

218

0

218

218

218

Total Flow

122,728

17

122,711

122,711

122,711

Total Enthalpy Flow

-356,727

-10,382

-350,498

-35,0498

-35,0498

Table 11 and Table 12 show the stream tables and equipment table for the gasification section of the bioethanol plant. Component and total flows are per mass basis with units of kg/hr while total enthalpy flows are in units of MJ/hr. Gasification section stream table

47


Table 11. Stream table of process streams entering and exiting gasification section of the bioethanol plant.

STEAM

MILLED FEED

SYNGAS

O2-H2O

C

0

55,689

0

0

H2

0

523

10112

0

N2

0

604

0

0

O2

0

64,776

0

64776

CH4

0

0

355

0

CO

0

0

121799

0

CO2

0

0

11709

0

H2O

93,992

11

0

6070

NH3

0

0

735

0

HCL

0

0

812

0

H2S

0

0

119

0

CL2

0

789

0

0

S

0

101

0

0

Solids

0

218

218

0

93,992

122,711

145,859

70846

-1,099,100

350,498

-342,170

-16,262

Component

Total Flow Total Enthalpy Flow

48


Table 12. Material and energy inflow and outflow to the equipment of the gasification section of the bioethanol plant. R-101 [GASIFIER]

Component

IN

OUT

C

0

55,689

0

0

H2

0

523

10,112

0

N2

0

604

0

0

O2

0

64,776

0

64,776

CH4

0

0

355

0

CO

0

0

121,799

0

CO2

0

0

11,709

0

H2O

93,992

11

0

6,070

NH3

0

0

735

0

HCL

0

0

812

0

H2S

0

0

119

0

CL2

0

789

0

0

S

0

101

0

0

Solids

0

218

218

0

93,992

122,711

145,859

70,846

-1,099,100

350,498

-342,170

-16,262


Table 13 and Table 14 show the stream tables and equipment table for the gas cleaning section of the bioethanol plant. All flows are per mass basis with units of kg/hr while total enthalpy flows are in units of MJ/hr. 49


Table 13. Stream table of process streams entering and exiting the gas cleaning section of the bioethanol plant.

Component

SYNGAS

CYCLONE

CYCLONE

OVERFLOW

UNDERFLOW

PRECOOLED SYNGAS

CLEAN

H2S

COOLED

FERMENTER

NH3

SYNGAS

REMOVED

SYNGAS

FEED

REMOVAL

H2

10,112

10,112

0

10,112

10,112

0

10,112

10,112

0

CH4

355

355

0

355

355

0

355

355

0

CO

121,799

121,799

0

121,799

121,799

0

121,799

121,799

0

CO2

11,709

11,709

0

11,709

11,709

0

11,709

11,709

0

NH3

735

735

0

735

735

0

735

0

735

HCL

812

812

0

812

812

0

812

812

0

H2S

119

119

0

119

119

119

0

0

0

Solids

218

0

218

0

0

0

0

0

0

Total Flow

145,859

145,641

218

145,641

145,641

119

145,522

144,906

735

-342,170

-560,080

217, 910

-654,220

-436,300

-1.7

-587,620

-585,660

-1,963

Total Enthalpy Flow

50


Table 14. Material and energy inflow and outflow to the equipment of the gasification section of the bioethanol plant. GC-101 [CYCLONE]

Component

H2

IN

OUT

E-101

GC-102 A/B

E-102

GC-103

[Heat exchanger]

[Adsorption column]

[Heat exchanger]

[Ammonia scrubber]

IN

OUT

IN

OUT

OUT

IN

OUT

IN

OUT

OUT

10,112

10,112

0

10,112

10,112

10,112

0

10,112

10,112

10,112

10,112

10,112

0

CH4

355

355

0

355

355

355

0

355

355

355

355

355

0

CO

121,799

121,799

0

121,799

121,799

121,799

0

121,799

121,799

121,799

121,799

121,799

0

CO2

11,709

11,709

0

11,709

11,709

11,709

0

11,709

11,709

11,709

11,709

11,709

0

NH3

735

735

0

735

735

735

0

735

735

735

735

0

735

HCL

812

812

0

812

812

812

0

812

812

812

812

812

0

H2S

119

119

0

119

119

119

119

0

0

0

0

0

145,859

145,641

218

145,641

145,641

145,641

119

145,522

145,522

145,522

144,787

-336,600

-560,080

-217,910

-560,080

-645,220

-645,220

-1.7

77,605


145,522

-

-

-

-

436,300

587,620

587,620

585,660

Table 15 and Table 16 shows the stream table and equipment table for the fermentation section of the bioethanol plant. Component and total flows are per mass basis with units of kg/hr while total enthalpy flows are in units of MJ/hr.

51

0 -1,963


Table 15. Stream table of process streams entering and exiting the fermentation fermentation section of the bioethanol bioethanol plant. COOLED

PUMP

MIXED

MIXED

BROTH

BROTH

0

0

0

0

0

0

11,709

0

H2O

0

NH3 HCL

FERMENTER

FRESH

RECYLED

MIXED

FEED

BROTH

BROTH

BROTH

10,112

0

0

CH4

355

0

CO

121,799

CO2

Component

H2

FERMENTER

DC-1

PRODUCT

FEED

0

0

0

315

0

0

0

0

355

0

0

0

0

0

0

0

0

0

0

0

0

6188

31,040

56,531

87,571

87,571

87,571

93,274

93,274

308

0

0

0

0

0

0

0

0

0

812

0

0

0

0

0

0

0

812

0

0

0

0

0

0

0

0

H2S

WASTE

Ethanol

0

0

6

6

6

6

9,509

9,509

93

Acetic Acid

0

0

8,467

8,467

8,467

8,467

10,398

10,3 98

119,579

Total Flow

144,787

31,040

65,003

96,043

96,043

96,043

113,181

113,181

127,650

-585,660

-495,460

-942,330

-1,437,800

-1,455,800

-1,455,800

-1,614,500

Total Enthalpy Flow

52

-1,614,500

-61,705


Table 16. Material and energy inflow and outflow to the equipment of the fermentation section of the bioethanol plant. R-102 [Fermentation vessel]

Component

IN

E-103 [Heat exchanger]

OUT

IN

P-101 A/B [PUMP]

OUT

IN

P-102 A/B [PUMP]

OUT

IN

OUT

H2 [kg/hr]

10,112

0

0

315

0

0

0

0

0

0

CH4

355

0

0

355

0

0

0

0

0

0

CO

121,799

0

0

0

0

0

0

0

0

0

CO2

11,709

0

0

6,188

0

0

0

0

0

0

H2O

0

87,571

93,274

308

87,571

87,571

87,571

87,571

93,274

93,274

NH3

0

0

0

0

0

0

0

0

0

0

HCL

812

0

0

812

0

0

0

0

0

0

0

0

0

0

0

0

0

0

0

H2S Ethanol

0

6

9,509

93

6

6

6

6

9,509

9,509

Acetic Acid

0

8,467

10,398

119,579

8,467

8,467

8,467

8,467

10,398

10,398

Total Flow [kg/hr]

144,787

96,043

113,181

127,650

96,043

96,043

96,043

96,043

113,181

113,181

Total Enthalpy

-

Flow [MJ/hr]

585,660

-1,437,800

-1,614,4500

-61,705

-1,437,800

-1,455,800

-1,455,800

-1,455,800

-1,614,500

-1,614 ,500

Table 17 and Table 18 shows the stream table and equipment table for the separation section of the bioethanol plant. Component and total flows are per mass basis with units of kg/hr while total enthalpy flows are in units of MJ/hr.

53


Table 17. Stream table of process streams entering and exiting the separation section of the bioethanol plant. DC-1

DC-1

FEED

BOTTOM

H2O

93,274

113,062

Ethanol

9,509

Acetic Acid Total Flow

Component

Total Enthalpy Flow

Component

DC-1

RECYCLED

Boilup

BROTH

113,062

56,531

56,531

40,417

40,417

40,417

3674.3

36,743

`12

`12

6

6

10,452

10,452

10,452

950.2

9,502

10,398

16,934

16,934

8,467

8,467

2,125

2,125

2,125

193.2

1,932

113,181

130,008

130,008

65,003

65,003

52,994

52,994

52,994

4817.7

48,177

-942,330

-942,330

-64,372

-643,720

HEATED BOTTOM

-1,614,500

DC-2 BOTTOM

DC-2 HEATED BOTTOM

DC-2

DC-2

Boilup

BOTTOM

DC-2 TOP

DC-1 TOP

DC-2 TOP CONDENS

DC-1 TOP

DC-1

DC-1

CONDENS

DC-2 VESSEL EXIT

VESSEL EXIT

DC-1 REFLUX

DC-2

DC-2

REFLUX

DISTILLATE

DC-1 DISTILLATE

ETHANOL

WWT

H2O

60,844

60,844

25,053

35,791

1,736

1,736

1,736

784

952

0

952

Ethanol

1,161

1,161

478

683

15,752

15,752

15,752

6933

8 ,819

8,819

0

Acetic Acid

3,284

3,284

1,352

1,932

1.6

1.6

1.6

.07

.09

0.09

0

Total Flow

65,289

65,289

26,884

38,406

17,405

17,405

17,405

7,634

9,771

8,819

952

-281,863

-402,661

-52,412

-66,673

-52,071

Total Enthalpy Flow

54

15,014


Table 18. Material and energy inflow and outflow to the equipment of the separation section of the bioethanol plant. T-101 Component

IN

E-104

OUT

P-104

E-105

IN

OUT

IN

OUT

IN

P-105

OUT

IN

OUT

H2O

93,274

56,531

36,743

113,062

113,062

113,062

113,062

40,417

40,417

40,417

40,417

Ethanol

9,509

6

9,502

`12

`12

`12

`12

10,452

10,452

10,452

10,452

Acetic Acid

10,398

8,46 7

1,932

16,934

16,934

16,934

16,934

2,125

2,125

2,125

2,125

Total Flow

113,181

65,003

48,177

130,008

130,008

130,008

130,008

52,994

5 2,994

52,994

52,994

-1,614,500

-942,330

-643,720

Total Enthalpy Flow

T-102 Component

IN

E-106

OUT

P-106

E-107

IN

OUT

IN

OUT

IN

P-107

OUT

IN

OUT

H2O

36,743

35,791

952

60,844

60,844

60,844

60,844

1,736

1,736

1,736

1,736

Ethanol

9,502

683

8,819

1,161

1,161

1,1 61

1,161

15,752

15,752

15,752

15,752

Acetic Acid

1,932

1,932

.09

3,284

3,284

3,284

3,284

1.6

1.6

1.6

1.6

Total Flow

48,177

38,406

9,771

65,289

65,289

65,289

65,289

17,405

17,405

17,405

17,405

-643,720

-575,230

-66,673

Total Enthalpy Flow

6.3.2 Utilities Side

Table 19 and Table 20 show the stream table and equipment table respectively for the utility streams of the entire bioethanol plant section of. Comp onent and total flows are per mass basis with units of kg/hr while total enthalpy flows are in units of MJ/hr. 55


Table 19. Stream table of utilities used in the entire bioethanol plant. Component

H20

DRYER-

DRYER-

CW1

LPS1

CW2

LPS2

CW3

STEAM

STEAM EXIT

69,945

CW3 RETURN

69,945

153,766

153,766

89,667

89,667

205,641

205,641

-73,485

-69,302

-2,451,400

-2,356,900

-1,427,900

-1,384,700

-3,278,400

-3,260,300

CW4

CWR4

DIST1-CW

DIST1-CW

DIST2-CW

DIST2-CW

Total Enthalpy Flow Component

RETURN

H20

DIST1-LPS

RETURN

DIST1-LPS

DIST2-

RETURN

LPS

6,469,180

6,469,180

1,490,110

1,490,110

827,897

827,897

60,645

60,645

27,262

103,130,000

102,570,000

23,737,000

23,633,731

13,188,000

13,130,424

937,570

1,069,866

421,480

DIST2-

P-101

P-102

P-104

P-105

P-106

P-107

MILLING

LPS

ELECTRICITY

ELECTRICITY

ELECTRICITY

ELECTRICITY

ELECTRICITY

ELECTRICTY

ELECTRICITY

8x10-4

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

.0253

Total Enthalpy Flow Component

RETURN

H20

827,897

Total Enthalpy Flow Electricity

-480,871

56


Table 20. Utilities side material and energy inflow and outflow to the equipment of the entire bioethanol plant Component

H20

PT-101

PT-102

E-101

E-102

E-103

69,945

69,945

153,766

153,766

89,666

89,666

205,641

205,641

73,485

69,302

-2,451,400

-2,356,900

-1,427,900

-1,384,700

-3,278,400

-3,260,300

Total Enthalpy Flow Electricity

.0253

Component

H20

E-104

E-105

E-106

E-107

P-101

56,531

56,531

1,490,110

1,490,110

27,262

27,262

827,897

827,897

-937,570

-1,069,866

-23,737,000

-23,633,731

-421,480

-480,871

-13,188,000

-13,130,424

Total Enthalpy Flow 8x10-4

Electricity Component

Electricity

P-102

P-103

P-104

P-105

P-106

P-107

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

57

7. Process Control 7.1 Control Overview This section of the report covers all the controls that are added to P&ID. The controls are arranged in the order in which they are presented in the P&ID. There are controls on the same type used more than once in the system and will be described only once here. Starting from the gasifier unit (R-101), we have steam and feedstock coming in at a specified ratio. Now we know that flow of steam and air can fluctuate and therefore introduce a disturbance into our system. Therefore we have implemented a ratio control for this section, in order to keep the ratio of feed to steam constant, entering the gasifier unit. Figure 1.0 displays how this ratio control is applied to the system. For many processes, a key objective is to maintain the flow rates of two process steams in some proportion to one another. In such cases, ratio control is applied. When ratio control is applied, one process input, the dependent input, is proportioned to the other process input, known as the independent input. The independent input may be a process measurement or its set point. The proportion that needs to be maintained is between the two inputs is known as the ratio. In Figure 9, the independent input measurement is the flow rate of feed coming into the gasifier. The ratio controller sets the set points of the flow controller rather than the valve position, as illustrated in Figure 9.

Figure 9: Ratio control loop design for steam to feed ratio


Thus any nonlinearity installed characteristics associated with the valves is addressed by the flow controllers and has no impact on the ratio controller being able to maintain the ratio set point. The input to the ratio control is the measurement or set point of the independent flow which is the flow of feed coming in to the gasifier. The ratio controller, multiplies this measurement to the ratio, to determine the set point of the dependent flow which is steam in our case. The set point of dependent flow will be sent to the flow controller that will manipulate valve v-6 shown in Figure 9 in order to keep the ratio constant. The measurement of both flows must be done as close to the gasifier unit to avoid any time delay in the response. Therefore as shown in Figure 9 the measurement of flow is taken right before the streams enter the gasifier unit. Controller type PI will be used in this situation because they eliminate forced oscillations and steady error resulting in operation of on-off controller and P controller respectively. However, introducing integral mode has a negative effect on speed of the response and overall stability of the system. Since in our case we are using a flow controller to manipulate the valve position, we know that this type of control has a very fast response time, therefore we don’t have to worry about the delay in response introduced by the integral mode. Furthermore, integral mode will eliminate any off-set that is present in the system. There are other section of the P&ID where ratio control is implemented. This included controlling the ratio between Feed entering each distillation column and the reboiler utility. Sometimes in the process we might decide to increase our production, so once we do that we need to change the reboiler duty as well. Therefore instead of a person going and manually adjusting the utility flow of the reboiler, we will use a ratio control to keep the flow of steam constant with the coming feed in the column. Furthermore a disturbance can occur and alter the flow of steam by either decreasing it or increasing it. This disturbance will also be eliminated by the ratio control strategy shown in Figure 10. Note that we have two distillation columns in our process and only the first one has this type of control strategy. This is because we are already using the feed entering the second distillation column to control the level of the reflux drum of the first distillation column.

59


Therefore we can’t have two controllers fighting and didn’t apply this strategy to the second column.

Figure 10: Ratio control design for Feed to reboiler utility ratio

The idea is similar to what was described in the previous ratio controlled system. In this scenario, the independent variable for ratio control is the measurement of the feed entering the column. This measurement is sent to the ratio control that multiplies it to the ratio and sends an output to the flow control which adjusts the position of valve V-496, as shown in Figure 10. The dependent variable in this case is the flow of LPS entering the reboiler. Similar to the previous case, flow measurements are taken right before the streams enter their desired unit. Measurement of feed flow is taken before it enters the distillation column and measurement of LPS flow is taken before it enters the reboiler. All this is done to avoid any type of time delay in the response of the control system. A PI controller type is suitable for this control system because as mentioned earlier they will eliminate forced oscillations and steady state error resulting in operation of on-off controller and P controller selectively. A key point to note which wasn’t mentioned earlier is that PI 60


controller does not increase the speed of response. It can be expected since PI controller does not have any means to predict what will happen with the error in the near future. This problem can be solved by introducing the derivative mode which has ability to predict what will happen with the error in the near future and thus to decrease a reaction time of the controller. From our prior knowledge of control theory, flow controllers have a very fast response time and therefore we have concluded just to use a PI controller for the ratio control applied to this distillation column. Temperature inside a distillation column is one of the variables that needs to be controlled. This is because distillation is temperature dependent; any variation in temperature will cause the purity of the product stream to decrease. A ratio control structure for this system is shown in Figure 11.

Figure 11: Ratio control structure between distillate and reflux to maintain a steady t emperature inside the column

The independent measurement for the ratio control is the flow of distillate leaving the reflux drum. This measurement is sent to a ratio control that multiplies it to the ratio. However another independent measurement is sent to the ratio control and this is the 61


temperature inside the column. The ratio control compares the set point of temperature to the set point stored by the operator and takes action by manipulating valve V-105 which either decreases or increases the reflux flow depending on what is the situation. The flow of reflux is measured right before it enters the distillation column and the flow of distillate is measured right after the stream leaving reflux drum (V-102) splits. Since this temperature control will affect the purity of the product stream, we want the temperature measurement to be taken from the top trays of the distillation column. A similar approach is used to determine the location of temperature measurement for the 2nd distillation column, it will be measured from one of the trays at the top. A PID controller is used for this control structure. They have all the necessary dynamics including fast reaction on change on controller input (D mode), increase in control signal to lead error towards zero (I mode) and suitable action inside control error are to eliminate oscillations (P mode). The reason behind using a PID controller is that this is the most important area in our system and it effects the purity of our desired product, therefore we want the control system to be perfect in all aspects mentioned above. The derivative mode improves the stability of the system and enables increase in gain K and decrease in integral time constant Ti, which increases speed of the controller response. From our prior knowledge or process control, we know that temperature control in systems have a slower response time, therefore we need to a controller type that is fast and will not change the temperature of the system and therefore keep the purity constant. Note that this type of control structure is applied to the second distillation column as well and therefore we haven’t shown it again as all the parts are similar to what is described above. Some of the units in the process require pressure control inside. This is because high pressure can lead to explosion of the unit and therefore cause damage to the surrounding units and also might kill workers around that area. Gasifier unit is one of the reactors that has a pressure control used in it. Figure 12 shows the control loop structure for pressure control in the gasifier unit (R-101)

62


Figure 12: Pressure control inside the gasifier unit

This is a regular single loop feedback control system which has a control variable and a manipulated variable. The controlled variable is the pressure inside the unit and the manipulated variable is the flow of the product stream leaving the gasifier unit. When the pressure inside the unit is too high, valve V-2 will open to push more vapour out the unit in order to decrease the pressure. When the pressure is too low, the valve will close slightly to keep the vapour inside the unit so that the pressure can reach its desired set point. Since gas is being formed in this reactor, the pressure measurement can be taken anywhere on the top reactor. However the measurement needs to be taken away from the outlet stream in order to avoid any errors in the reading. A PI controller will be a suitable type of controller for this system. Since pressure control by themselves have a fast reaction time, therefore we don’t need a derivative mode in this situation. We still require an integral mode to remove the offset and a proportional mode to eliminate any oscillations. Another section of the system where pressure control is used are two distillation columns. Both the columns have a similar structure of the pressure control system and therefore only one is explained in detail here. Figure 13 shows the structure of pressure control loop designed on the distillation column. 63


Figure 13: Pressure control loop design for the distillation columns

This type of pressure control strategy is similar to that which was used in the Gasifier unit (R-101). However this one has a different manipulated variable. In order to control the pressure inside the column, we are manipulating the cooling water flow entering the condenser. The idea is to reduce vapour accumulation at the top of the distillation column during high pressure scenarios. In a high pressure scenario, the position of valve V-83 in Figure 13 will open more to let more cooling water enter the condenser. This eventually will liquefy more vapour and therefore will decrease the pressure inside the column. The pressure sensor must be placed somewhere on top of the column away from the exit location to avoid any errors in the measurement. A PID controller type will be suitable for this system. Even though we have mentioned earlier that pressure control is fast in terms of dynamics but in this scenario we time delay. When the cooling water flow will increase, it will take time for the vapour to condense in the condenser and therefore it will take time for the pressure reading to 64


change inside the column. So we want a controller that has fast response and can predict the future error using the derivative mode. Sometimes when we are dealing with liquid systems, we have a problem of flooding in the vessel. In order to avoid this situation, level control is applied to different vessels that are used in the system. One of this includes level control inside the fermenter unit (R-102). Since the product is liquid leaving the fermenter, we do not want the liquid to fill up the fermenter and therefore flood it. A simple level control structure for this system is shown in Figure 14. Level control is also a single loop feedback control that uses a control variable and a manipulated variable. The control variable in this case is the level of liquid inside the fermenter and the manipulated variable is the valve position of V-550, which changes the flow of product stream leaving the fermenter.

Figure14: Level control structure for the fermenter

This type of level sensor calculated the hydrostatic pressure inside the unit at two different heights and the difference gives us the level of liquid in the tank. The location of these hydrostatic sensors is determined by what is the maximum allowable level that can lead to safe operation. The minimum level is determined by how fast we are pumping the 65


liquid out. Since pumps cannot run dry, some level of liquid needs to be present all the time in order to avoid dry operation of the pump. If the level inside is high, the level controller will send a signal to valve V-550 which will open more in order to reduce the level inside the fermenter and bring it back to the desired set point given by the operator. A PI control type is suitable for this situation. Since level controls have a fast response time and the hydrostatic measurement itself is automated, therefore applying a derivate mode here won’t make a difference in terms of improving the response of the system. Reflux drum vessels used to store the liquid after the condenser in a distillation column also requires level control to avoid flooding of the vessel. Figure 15 shows the control loop structure design of the level control used for V-101. Note that the reflux drum for the second distillation column has a similar control strategy being applied and therefore is not mentioned in detail.

Figure 15: Level control for reflux drum

66


The level control here uses a similar strategy as the previous level control system. Two hydrostatic pressure measurements are taken from the vessel and the corresponding output is sent to a level controller which manipulates the position of valve V-546. This type of level control uses a single feedback control loop designed with a PI type controller. As mentioned above that we don’t need a derivative mode since we are not looking for an improvement in the control system, its already operating at optimum conditions. From prior knowledge we can conclude that this type of response behaviour between level and flow is stable and therefore further stability in the system is not required by the input of derivative mode. When a heat exchanger is used to heat up a desired stream or cool it down. Temperature of the product stream leaving the heat exchanger must be controlled. This is because any fluctuations in the flow or temperature of the utility stream can cause deviation in the temperature of the product stream exiting the heat exchanger and we might not get the desired temperature output that we are looking for in the product stream. Therefore a cascaded temperature control strategy is applied to all the heat exchangers and condenser used in the process and only one of them is explained in detail here. Figure 16 shows how one these heat exchangers have a cascaded temperature control being implemented.

67


Figure 16: Cascaded temperature control design around the condenser E-104

A certain degree of flexibility needs to be added around the heat exchanger and condensers, since the flow rate of the CW coming in might fluctuate and act as a disturbance to the temperature of the product stream leaving the exchanger. To avoid this disturbance, a cascade loop is implemented around the condenser as shown in Figure 16. The inner loop of this cascade control is measuring the flow of the CW and controlling it by manipulating the pneumatic valve V-68. This valve is also labelled fail closed because during a failure, if the valve is in the closed position the utility is not wasted. The outer cascade loop measures the temperature of the product stream exiting the condenser, and this is the set point for the inner cascade loop. Therefore 2 controllers are required for keeping the temperature of the product stream constant when a disturbance occurs in the flow of CW. The good thing about having a cascade control is that the inner loop will have a much faster dynamic response than the outer loop, therefore a disturbance will have a minimal effect on the temperature of the product stream. The temperature sensor which measures the temperature of the product stream exiting the heat exchanger is

68


located right after the stream exits the heat exchanger to avoid the disturbance in reading from any temperature losses along the pipe. The flow sensor for the measurement of CW flow is located before the pneumatic valve V-68. The sensor could also be located after the pneumatic valve but since the pressure in the change will change on both sides of the valve, locating the flow sensor on either side will not have a drastic effect on the control strategy being applied. In other words it doesn’t matter where we put the flow sensor, we are going to get the same control behaviour. However we cannot have the sensor far away from the heat exchanger, since there will be error in the reading due to pressure losses along the pipe. A PID type controller is best suitable for this control scheme. This is because temperature change will take time once the flow rate of the utility is increased therefore the system dynamics will be slow giving rise to a very high time delay in the control response. The derivative mode will decrease the integral time constant Ti as mentioned earlier, and will therefore increase the speed of the controller response. A key point to note here about derivative mode that wasn’t mentioned in the previous sections, is that it’s not taken from the error signal but rather from the system output variable. This is done to avoid effects of the sudden change in the value of error signal. Sudden change in error signal will cause sudden change in control output. To avoid that it is suitable to design D mode to be proportional to the change of the output variable. pH is another parameter that needs to be controlled in a system where there is a pH sensitive medium. In our case, the bacteria in the fermenter works best at a pH of 6, therefore it is our goal to keep the pH of the fermenter constant at 6 for optimum conditions. A cascade control structure is applied for this pH control, Figure 17 shows how the control loop is designed around the fermenter.

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Figure 17: pH control loop structure for the fermenter.

The reason behind using a cascade control for this system is that pH control itself is very slow in terms of dynamics of the system. We are using a single pH probe to measure the pH of a huge fermenter, it will take a lot of time for the pH of the fermenter to change, and therefore we need a very fast control system for this case. The pH measurement is sent to a flow controller which manipulates the position of valve V-78. The inner loop which controls the flow has a much faster response than the outer loop which controls the pH and therefore the inner loop will run much faster than the outer loop, keeping the pH constant at the desired set point. This type of cascade control will involve using a PID control. Since we already have a slow response time in the system, we need to have a PID type controller so that the D mode improves the stability and increases the speed of the controller response. The pH measurement needs to be taken at the bottom of the liquid level, away from any mixer or inlet or exit. I have previously done a pH control lab and the problem in that was the location at which pH measurement was taken. Therefore it’s very important that the pH probe is located in the area where there are no fluctuations in flow and we have a steady flow profile. Since we are recycling our broth back into the fermenter, we need to purge some of it out of the system to avoid any accumulation inside our units. The purge stream needs 70


to be in a certain ratio with the recycle stream, this is achieved by having a ratio control as shown in Figure 18. Since the flow of recycle stream can change and act as a disturbance, we want to fix our flow of purge so that we don’t remove extra stuff from our system and save cost.

Figure 18: Ratio control structure between purge stream and recycle stream

As mentioned in the earlier sections, ratio control involves independent and dependent variables. The independent variable is the flow of recycle stream and the dependent variable is the flow of the purge stream. The independent measurement of the recycle flow is sent to the ratio controller that multiplies it by the ratio and manipulates the position of valve V-562 accordingly. The flow sensor of the recycle stream is attached right before the split is made to avoid any error is measurement that can be caused by pressure losses in the pipe. Similarly, the flow measurement for purge stream is done right after the spilt is made to avoid any errors due to pressure losses. Using a PI controller alone without the derivative mode will be suitable for this case since we are only doing a ratio control with flow which has fast control dynamics. Lastly we need to control the pH inside the unit S-101. Since everything exiting S-101 enters our fermenter, we don’t wait to disturb the pH of the fermenter and therefore we

71


need to keep S-101 at out desired pH level of 6. Figure 19 demonstrates how this type of control structure is designed around the unit S-101. This type of control design also involves a ratio control between recycle stream and the fresh medium stream entering the unit S-101. The independent measurement will be the flow rate of recycle stream and the dependent measurement will be the flow rate of the fresh medium stream. Since the flow rate and pH of the fresh medium stream can change and therefore can act as a source of disturbance in our system, it needs to be controlled.

Figure 19: pH control design for unit S-101

The pH measurement from the unit S-101 is sent to the ratio controller which compares it to the set point of pH that we have defined already in the system, along with 72


this input the ratio control also gets an input from flow of the recycle and accordingly adjusts the flow of the fresh medium entering the unit. The location of the flow sensor for the recycle is done right before the mixing point so that we can avoid errors due to pressure drop in the pipes. Similarly, the location of the flow sensor for the fresh medium is right before the mixing point to avoid any errors in flow measurement due to pressure losses in pipe that can lead to drop in the flow across the pipe. In order to get an accurate pH measurement inside the unit S-101, the probe needs to be fully emerged in the liquid medium all the way till the bottom, away from any inlet and outlet, to avoid any error in pH measurement due to flow fluctuations. We need to use a PID controller for this section because as mentioned earlier pH system dynamics are very slow and we need a derivative mode for fast response of the controller. The derivative mode will look at the slope of the error and decide what action to take. In other words we will have a feedforward control strategy being applied by looking into the future of our error and predicting what it’s going to be so that we can take the appropriate action in the present. This brings us to the end of all the control loops that are used in the system. Now the final copy of all the sections of the P&ID will be presented that consist of all the control loops that were mentioned above. It also contains those control loops that were mentioned earlier but not explained in detail because they had the same control design as those which were mentioned.

7.2 Preliminary P&ID of Process The process is split into different sections so that the P&ID can fit into this document. The first section which doesn’t require any controlling is the pre-treatment section shown in Figure 20.

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Figure 20: Pre-treatment section

The output of the pre-treatment goes into the gasification unit which is described in the next section. The second section of the P&ID consists of the gasifier unit and all the control loops around it demonstrated by Figure 21.

Figure 21: Gasification section

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The output of the Gasification unit goes into a series of gas cleaning steps which are shown in Figure 22.

Figure 22: Gas cleaning section of the P&ID

The next section of the P&ID displays the fermenter unit and the control’s that are applied and the recycle storage unit S-101 with the pH control and other controls are also shown.

Figure 23: Fermenter section of the P&ID

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The final section of the P&ID consists of the distillation and is shown in Figure 16.

Figure 24: Summary of the Distillation section of the process

8. Equipment design, sizing and costing – process side 8.1 Costing overview Moving on from the P&ID and process design discussion, this section aims to present a study estimate of the costs of the plant, though the costs presented are not bare module costs so they do not include installation factors. Because this is a study estimate, there are large margins of error associated with each capital cost, some of which can be up to 200%. Additionally, both capital costs and operating costs use results from the Aspen Plus simulation, which were presented previously in stream tables. These results are not exact and may contain error due to the difficulty in accurately modeling the gasification and continuous fermentation processes as well as other small differences between the simulation or model and reality. The margin of error for all costs is assumed to be ± 50%. 76


Capital costs were calculated using cost correlations in (Seider et al., 2009) (Towler and Sinnott, 2012) and (Woods, 1983). The costing method used for both capital costs and operating costs are explained in their respective sections. Additional important details regarding the economics of the plant are listed here: 

The MARR was chosen as 10%.



The plant lifetime was chosen as 25 years.



Nominally, the plant runs for 24 hours and 330 days a year but since this is unrealistic due to extra time required on some days for maintenance, the hours lost from not running the plant can be made up by adding more days of operation per year.



The first year is used to buy, setup, test and run the plant at lower capacity in order to troubleshoot any problems that come up during initial operation. The plant starts operating in the second year.



The tax rate is 25%.



Plant economics are initially analyzed without an ethanol subsidy or a carbon tax.



Transportation costs are ignored completely for this analysis due to the inability to accurately document which population centers are using the fuel, what their demand is and what the travel pathways are in order to reach these destinations.



The following USD to CAD conversion rate was used throughout: 1 USD = 1.25 CAD.

8.2 Capital costs Capital costs were calculated for the following components: dryer, crusher (hammer mill), gasifier, cyclone, heat exchangers, fermenter, vessels, pumps, distillation columns and storage tanks. In addition, the only component which was cost as working capital was zinc oxide, which is the adsorbent used in the H2S gas cleaning adsorber column. All costs were converted to 2013 dollars using CEPCI inflation factors. The final costs were converted to Canadian dollars from U.S. dollars, so the entire economics analysis is done in Canadian dollars. While Woods’ textbook uses cost information and correlations from the 1970s, these correlations were only used for vessels, pumps and 77


distillation column costs. However, the fact that distillation columns are a significant portion of the final capital cost makes this final estimate a weak one as cost correlations from the 1970s cannot be used with confidence in modern times. The other two textbooks use cost information and correlations from 2006 so they are far more accurate. During cost estimation, some of the cost correlation factors exceeded their given bounds. Although this means that the associated capital costs are not as accurate, these deviations are not significant relative to the existing error associated with each cost (approximated as 50%), especially since the bounds are not exceeded by much. The dryer was cost as a spray dryer because it uses the evaporation rate in the correlation, which is one of the design factors chosen during the dryer design. The crusher was cost as a hammer mill as this is the type of crusher chosen in the plant design. Since there is no direct gasifier cost correlation in the books used, the gasifier was cost as a pyrolysis furnace. This models the gasifier accurately because one of the key steps in the gasification process is pyrolysis and the gasifier essentially takes on the form of a furnace as miscanthus is fed and burned. For gas cleaning, the cyclone was cost using the gas flow rate correlation, and the H2S adsorber and wet scrubber were cost as vessels using a volume correlation. The volume for the adsorption vessel was calculated by equating it to the volume of the adsorbent during operation, which used the vessel’s adsorption capacity and density and an operation time of 15 days. An operation time of 15 days was picked to limit the size of the adsorption column. Heat exchangers were cost using the standard heat exchanger area correlation. This heat exchanger area was calculated during heat exchanger design. Specifically, the closest overall heat transfer coefficient was chosen to suit the heat exchanger’s shell and tube species based on knowledge and experience, then heat exchanger duty was found from the Aspen Plus simulation using the HEATER model and finally the corresponding area was calculated. Since there was only one fermenter, there was some freedom regarding the design specifications and type of tank to use. The final chosen design was a jacketed, closed-lid stirred tank design. As a result, the cost correlation used fit perfectly here as it was for a

78


jacketed, agitated design. However, the fermenter was far too big and the cost factor exceeded its bounds by far too much. Therefore, the fermenter was cost as 10 fermenters and these costs were added together. There were two distillation column condenser vessels and both were cost using a Woods’ cost correlation involving the length and diameter. Pumps were cost using the standard power/electricity usage correlation. However, since the Aspen Plus simulation did not initially include pumps (as there was no pressure loss), several pumps were duplicated. These were duplicated from two pumps that were added to specific parts of the simulation or flow sheet where the flows differed, namely the recycle part and the fermentation outflow part. In the end, the combined capital cost of pumps was far lower than the other capital costs so this duplication of pumps did not greatly affect the final capital cost. The effect was considered ne gligible. Distillation columns were cost using a correlation in Woods based on diameter and tower height. The tower height was simply calculated by multiplying the tray spacing found from the Aspen Plus simulation sizing analyzer by the number of trays. However, the sizing analyzer gave a diameter far larger than the length, which does not resemble the shape of an actual distillation column and cannot be transported by truck. Instead, a 5 m diameter was chosen so that it could be transported but the volume required (based on feed flow rate) raised the height to >400 m. Since this is not feasible either, the distillation column in reality would have to be divided up into multiple distillation columns of height of around 50-80 m, as shown in the P&ID. Storage containers had their own correlation. The mixing tank that is before the fermenter in the P&ID needs to be a closed-lid tank that is kept at fermenter operating conditions so it was cost in the same way the fermenter was cost. As a result, its cost ended up being the same as that of the fermenter. Total capital costs for each component of the plant and the total capital cost for the entire plant is shown in Table 21.

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Tabl e 21 . Sum of capital costs for each type of unit and total capital cost.

Expense

CAD Cost

Pretreatment

$

2,104,000

Gasifier

$

8,182,000

Cyclone

$

729,000

Gas Cleaning

$

1,208,000

Heat Exchangers

$

1,598,000

Fermenter

$ 20,465,000

Vessels

$

516,000

Pumps

$

83,000

Distillation Columns $ 27,919,000 Storage

$ 21,953,000

Zinc oxide

$

TOTAL

2,374,000

$ 87,129,000

Table 21 shows that the largest costs are for the gasifier, fermenter, distillation columns and storage containers. Storage costs are inflated due to the pre-fermenter mixing tank being cost as a reactor instead of a storage tank. This was discussed in the previous paragraph. When comparing the largest costs, it is clear that the highest cost is that of the distillation columns. These distillation columns are unusually large in size due to the high feed flow rate. This is because there is no real parallelism in the process, as the fermenter exit contents are not split into multiple streams and instead go through two distillation columns in series. Although this was remedied by splitting the feed flow into multiple distillation columns, this does not lower the high cost of the columns. In fact, since multiple distillation columns require more material to be produced than a single long distillation column, splitting the feed as stated would actually increase the cost even more, possibly to around $30 million. However, this nonlinearity was not considered 80


when adding more columns and it was assumed that splitting the feed would not change the cost and not have an effect on the separation performance. The next largest cost, ignoring storage costs, is that of the fermenter. The fermenter, unlike the distillation columns, does not require feed splitting (even though it was cost as multiple fermentation tanks) because it can be extended in width and height without too much trouble. The only thing to consider here would be the gas-liquid mass transfer rate, contact area or collision rate, all referring to how much the gas and liquid mix in the tank. However, the syngas only takes a few minutes to contact and react with the contents in the fermenter and since the gas moves quickly, it is assumed that the edges of the tank are reached fairly easily by the syngas. Therefore, it is not as critical to split the feed into multiple fermenters to reduce the size. In addition, unlike the distillation columns, running the fermentation process is a lot more difficult in parallel due to the requirement of each fermenter to maintain operating conditions like pH, temperature, pressure and ethanol levels under strict ranges. This would also require additional safety systems, more operators, splitting the recycle stream running through the plant and more mixing tanks or storage containers. The third largest cost is the gasifier cost. It was initially assumed that this would be the largest cost due to its position in the flow sheet at the beginning where it is one of the first units to be in contact with the miscanthus feedstock. However, since the capital cost is simply cost as a pyrolysis furnace, the cost is low because the expected extra costs in maintenance of the gasifier and energy management are not taken into account here. Pretreatment, gas cleaning and heat exchangers make up the next three biggest capital costs. These costs are relatively proportional to their role in the syngas fermentation process. Gas cleaning costs (including the cyclone) add up to almost $2 million, which is ironically smaller than the cost of the zinc oxide adsorbent. Zinc oxide was cost by obtaining its volume through its adsorption capacity and density as stated in the previous paragraph (in the adsorption section). The pump capital cost is relatively the smallest cost in Table 21. This shows that having two pumps side by side (one as a backup) is not only feasible but is the best option

81


to take when building a syngas fermentation plant. For this reason, only two pumps out of the seven pumps were single (without a backup).

8.3 Operating costs Operating costs were calculated on an annual basis and like all other costs, they are in Canadian dollars. Utilities make up the largest portion of the operating requirements in terms of the role they play in running the plant and their constant need in order to keep the plant safe. The electricity price is the on-peak electricity price in Ontario: 14 cents per kilowatt-hour. This may not be reasonable for the current time but it is there to account for changes (most likely increases) in the electricity price over the plant lifetime, which will last multiple decades. In other words, it is there for contingency. All other utility prices were taken from (Seider et al., 2009) presented in the costing overview section. Utility operating costs were calculated by simply multiplying the cost factor by the utility flow over the annual duration of plant operation (24 hours for 330 days). The three other important operating costs are the costs of the feedstock, broth and wet scrubber operation. The feedstock cost is based on a cost factor estimate of $71.50/tonne for growing and harvesting the miscanthus in Canada (Roy, 2014). This includes costs to run an irrigation system, to properly plant and maintain the crops and the cost of the fertilizer. The wet scrubber annual cost is based on the flow of the waste stream according to a cost factor by the Environmental Protection Agency (EPA) in their Air Pollution Control Technology Fact Sheet (U.S. Environmental Protection Agency, n.d.) . The waste stream carries the separated impurities ammonia, hydrogen chloride and carbon dioxide at conditions close to standard temperature and pressure. The broth is cost as a fermentation medium known as corn steep liquor. Various journals, such as Maddipati, 2011 and Saxena, 2012, have recognized its low cost and effectiveness as a medium for the bacteria to produce ethanol in comparison to yeast extract, the traditional fermentation medium. It provides the necessary nutrients and conditions for the bacteria to thrive. The cost was calculated using by using a factor of $0.31/L of corn steep liquor but the entire stream of new broth was not cost in order to 82


counteract the overestimation of new broth required according to the Aspen Plus simulation. Other operating costs include the pretreatment steps, which require large amounts of energy input in the form of electricity for the crushing step and medium pressure steam for the drying step. These were added to the total operating cost. The final additions to the annual operating cost were the labour and infrastructure costs. Labour costs include the salaries and work related costs of operators, supervisors, engineers, maintenance personnel, management, and lab technicians. It was assumed that 10 operator posts would be needed to run the 10 primary processes, and that 4.4 people would be needed per post to run the plant 24/7 while changing shifts. Other infrastructure costs include overhead, maintenance materials, insurance, property tax, laboratories and operating overhead. Carbon tax was included but is only used when performing the sensitivity analysis. The complete operating costs are shown in Table 22. Tabl e 22 . All operating costs for the syngas fermentation plant.

Expense

Factor

Cost/Unit

Annual Cost [CAD]

Feedstock

$71.50/tonne

$

86,873,000

Wet Scrubber Annual Cost

$64000/waste stream

$

104,000

Raw Material (Broth)

$0.31/L

$

5,612,000

Electricity (P-101 A/B)

$0.14/kW-hr

$

1,811

Electricity (P-102 A)

$0.14/kW-hr

$

1,811


$0.14/kW-hr

$

1,171

Electricity (P-104 A)

$0.14/kW-hr

$

1,171


$0.14/kW-hr

$

1,171


$0.14/kW-hr

$

1,171


$0.14/kW-hr

$

1,171

Electricity (M-102)

$0.14/kW-hr

$

3,511,000

MPS to M-101

$10.5/1000 kg

$

7,271,000

LPS to E-103

$6.6/1000 kg

-$

6,085,000

LPS to E-106

$6.6/1000 kg

-$

4,078,000

CW to E-101

$0.020/m3

$

83

36,000


CW to E-102

$0.020/m3

$

21,000

CW to E-104

$0.020/m3

$

48,000

CW to E-105

$0.020/m3

$

295,000

CW to E-107

$0.020/m3

$

164,000

CW to FE-101

$0.020/m3

$

1,514,000

(A) Operators 1 post - 4.4 people

$70,000

$

3,080,000

(B) Supervision & Engineering

0.25*(A)

$100,000

$

770,000

(C) Maintenance Personnel

0.03*(fixed cost)

$75,000

$

2,091,000

(D) Engineering & Management

0.5*(A)

$100,000

$

1,540,000

Overhead

0.4*(A+B+C+D)

$

2,992,000

Maintenance Materials

0.03*(fixed cost)

$

2,091,000

Insurance

0.01*(fixed cost)

$

697,000

Property Tax

0.02*(fixed cost)

$

1,394,000

Laboratories

0.15*(A+B+C)

$

891,000 $

Carbon Tax Operating overhead

0.25*(A+B+C+D)

TOTAL

$

1,870,000

$ 112,713,000

Table 22 clearly lists the most costly operations. Disregarding labour, the largest operating costs, ones that exceed $1 million annually, are for feedstock, adding fresh broth to the fermenter, electricity and medium pressure steam for pretreatment, and cooling water for the fermenter. The feedstock cost exceeds all other operating costs by far but is not unusual by itself. It is large due to the various costs associated with growing the feedstock and maintaining it, which essentially requires as much work as it does to start and maintain a massive food crop year round. The two costs for the fermenter are expected since the fermentation process is at the heart of this process and is one of two big processes in the plant, the other being gasification. To maintain the health of the bacteria and the resultant ethanol production yield, high quality broth will always be required. As well, temperature control of the fermenter using cooling water is essential to maintaining the bacteria’s ability to produce

84


ethanol, as fermentation is exothermic and can inhibit the process. To summarize, both the cooling water and broth provide the ideal conditions to make this process work. Therefore, cost cuts to this operation are not recommended. The pretreatment or milling processes are both large costs, with the medium pressure steam going to the drying process being the most costly. In fact, this cost is the largest utility cost. The use of medium pressure steam instead of low pressure steam should be re-evaluated as this drives up the cost greatly. Electricity to the crushing process is needed, but other types of crushing processes (other than hammer mills) should be investigated to see if they can use less electricity or energy. The two operating costs shown for the low pressure steam going to the two distillation column reboilers are negative because they subtract the profit made from selling low pressure steam made in the heat exchangers. In reality, part of the low pressure steam made in the HX’s goes to the reboilers while the rest is sold but this still gives the same negative costs. Cost savings can be made on these two reboiler costs as well, which can be accomplished by reducing the boilup ratio of the bottoms exiting the distillation columns. However, this may reduce the purity of the final ethanol product, resulting in a less efficient process. Instead, it is recommended that the distillation columns should be further optimized or other changes should be made like feeding vapour and liquid from one stage to another, so that the reboilers are not used as much. In this way, the costs of the condenser utilities are brought down as well. As seen in Table 22, labour and infrastructure costs make up a large part of the operating costs and these costs are essential to running and maintaining the plant. However, it should be noted that these costs are based on fundamental assumptions like the amount of operators needed and rely on multiplication factors that can only give rough estimates. Compared to a total capital cost of almost $90 million, the total annual operating cost is relatively large as it is on another order of magnitude. Though this may indicate that the operating cost is too large, it is in fact the capital cost that is very low due to the simplicity and linear nature of the syngas fermentation process. Still, operating costs can

85


and should be brought down by saving on feedstock costs and making the process more efficient, especially by making it energy efficient.

8.4 NPV NPV calculations were performed using the previous total capital cost and total annual operating cost values. The costing overview section gives some of the key information used in these NPV calculations like the tax rate and the MARR. It is also important to note that the rate of depreciation on the equipment bought using the initial capital is 30% annually according to Class 43 of Canada’s CCA (capital cost allowance) law (Canada Revenue Agency, 2015). Since no other byproducts of the process are being sold except for excess steam, the only revenue that is made comes from selling the ethanol fuel. In Canada, ethanol can be used in multiple blends. These are E10, a 10% v/v (volume %) ethanol blend, or E85, an 85% v/v ethanol blend. These blends are only made by volume, and the rest of the blends contain gasoline. The sale of both these blends is not currently widespread in Canada and E85 is currently only commercially available in a select few areas. For this reason, US prices of ethanol were used as it is more widespread and the national average is more indicative of the price of ethanol. The E85 price was used from the US Department of Energy’s Alternative Fuels Data Center (U.S. Department of Energy, 2015). Since we know the blend according to volume percentage and prices are given by volume, the following formula (Equation 1) involving the price of E85 and gasoline was used to calculate the price of pure ethanol (E100).

$. $.   −.·  

  100 = 

.

86

  · . = $.

(1)


The process of syngas fermentation discussed in this report produces 99% ethanol so this price is appropriate given the purity. NPV calculations are shown in Figure 25 Period

All Income 0 $ 1 $ 90,000,207 2 $ 90,000,207 3 $ 90,000,207 4 $ 90,000,207 5 $ 90,000,207 6 $ 90,000,207 7 $ 90,000,207 8 $ 90,000,207 9 $ 90,000,207 10 $ 90,000,207 11 $ 90,000,207 12 $ 90,000,207 13 $ 90,000,207 14 $ 90,000,207 15 $ 90,000,207 16 $ 90,000,207 17 $ 90,000,207 18 $ 90,000,207 19 $ 90,000,207 20 $ 90,000,207 21 $ 90,000,207 22 $ 90,000,207 23 $ 90,000,207 24 $ 90,000,207

All eligible expenses $ $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359) $ (112,713,359)

All ineligible expenses Book value $ (87,129,306) $ 87,129,306 $ - $ 74,059,910 $ - $ 51,841,937 $ - $ 36,289,356 $ - $ 25,402,549 $ - $ 17,781,784 $ - $ 12,447,249 $ - $ 8,713,074 $ - $ 6,099,152 $ - $ 4,269,406 $ - $ 2,988,585 $ - $ 2,092,009 $ - $ 1,464,406 $ - $ 1,025,084 $ - $ 717,559 $ - $ 502,291 $ - $ 351,604 $ - $ 246,123 $ - $ 172,286 $ - $ 120,600 $ - $ 84,420 $ - $ 59,094 $ - $ 41,366 $ - $ 28,956 $ - $ 20,269

Depreciation $ 13,069,396 $ 22,217,973 $ 15,552,581 $ 10,886,807 $ 7,620,765 $ 5,334,535 $ 3,734,175 $ 2,613,922 $ 1,829,746 $ 1,280,822 $ 896,575 $ 627,603 $ 439,322 $ 307,525 $ 215,268 $ 150,687 $ 105,481 $ 73,837 $ 51,686 $ 36,180 $ 25,326 $ 17,728 $ 12,410 $ 8,687 $ 6,081

Taxable income Tax paid Net cash flow in period (TVM) NPV $ (13,069,396) $ (3,267,349) $ (83,861,957) $ $ (44,931,125) $ (11,232,781) $ (10,436,700) $ $ (38,265,733) $ (9,566,433) $ (10,865,057) $ $ (33,599,958) $ (8,399,990) $ (10,753,690) $ $ (30,333,916) $ (7,583,479) $ (10,333,770) $ $ (28,047,687) $ (7,011,922) $ (9,749,228) $ $ (26,447,326) $ (6,611,832) $ (9,088,775) $ $ (25,327,074) $ (6,331,768) $ (8,406,240) $ $ (24,542,897) $ (6,135,724) $ (7,733,492) $ $ (23,993,974) $ (5,998,493) $ (7,088,647) $ $ (23,609,727) $ (5,902,432) $ (6,481,260) $ $ (23,340,754) $ (5,835,189) $ (5,915,623) $ $ (23,152,474) $ (5,788,118) $ (5,392,837) $ $ (23,020,677) $ (5,755,169) $ (4,912,123) $ $ (22,928,419) $ (5,732,105) $ (4,471,640) $ $ (22,863,839) $ (5,715,960) $ (4,068,993) $ $ (22,818,633) $ (5,704,658) $ (3,701,544) $ $ (22,786,988) $ (5,696,747) $ (3,366,605) $ $ (22,764,837) $ (5,691,209) $ (3,061,546) $ $ (22,749,332) $ (5,687,333) $ (2,783,857) $ $ (22,738,478) $ (5,684,619) $ (2,531,183) $ $ (22,730,880) $ (5,682,720) $ (2,301,332) $ $ (22,725,561) $ (5,681,390) $ (2,092,283) $ $ (22,721,838) $ (5,680,460) $ (1,902,180) $ $ (22,719,232) $ (5,679,808) $ (1,729,320) $

Figure 25. NPV Analysis

Figure 25 shows unusual results as ethanol revenue exceeds the total capital but the yearly operating costs exceed both by a significant amount. Over the plant’s 25 year lifetime, NPV is negative and amounts to a loss of $223 million. This includes tax and depreciation as well as the yearly MARR. Therefore, the rate of return generated from this project does not exceed the MARR of 10% and makes this project not worthy of investment. It is also worth noting that a negative NPV of almost a quarter billion dollars is not much lower than the baseline of $0 NPV given the 25 year lifetime. (An NPV of $0 is the cut-off for project investment.) This means that the confidence level in rejecting investment into this project should not be high. Judging by the NPV values in Figure 25 across all the years, the NPV and the rate of return are reduced significantly by the high 87

(83,861,957) (94,298,658) (105,163,714) (115,917,405) (126,251,175) (136,000,403) (145,089,178) (153,495,418) (161,228,910) (168,317,557) (174,798,817) (180,714,440) (186,107,278) (191,019,401) (195,491,041) (199,560,034) (203,261,578) (206,628,183) (209,689,729) (212,473,586) (215,004,769) (217,306,101) (219,398,384) (221,300,564) (223,029,884)


operating costs, which take away from the revenues generated. This is primarily because of the high feedstock cost, as mentioned previously. It is therefore critical to lower operating costs by coming up with better ways grow feedstock while still maintaining the efficiency and yield from the process.

8.5 Sensitivity Analysis The sensitivity analysis was done as part of the previous NPV analysis. The following factors were varied to see possible changes in the NPV in different circumstances. 

Ethanol price



Feedstock cost (based on feedstock price)



Plant lifetime



Carbon tax

Ethanol price was varied instead of any kind of subsidy because a subsidy program which supports biofuel production already exists in Canada. It is known as The ecoENERGY for Biofuels Program and lasts until 2017 (National Resources Canada, 2014). Its incentive is based on the following formula (Equation 2) and the exact incentive amount is calculated further in the section.

()

Incentive = Eligible Sales L × Incentive Rate

  [2] $

L

Eligible Sales are assumed to be equal to the production capacity of ethanol from the plant. The Incentive Rate is a given rate that is fixed by the government of Canada for every year until 2017. Therefore, the ethanol price sensitivity shown here does not represent a subsidy but rather the normal variations in market price of ethanol. Although these normal variations change the NPV in a linear fashion in the sensitivity analysis, changes in the market demand due to price are also expected in reality. In addition, significant shifts in demand are expected due to the lower energy content of ethanol blends when compared to pure gasoline. Specifically, one gallon of E100 has around 73% of the energy content of gasoline (U.S. Department of Energy, 2014), so ethanol fuel will only be competitive in the market once the price of the ethanol is around 70% of the price of gasoline. For this to happen, the price would have to reduce drastically from current 88


levels. Therefore, a key assumption in this sensitivity analysis is that all the ethanol that is produced is sold at market price. In other words, there is market demand but it does not govern supply from this plant. The final part of the sensitivity analysis is for carbon tax. In British Columbia, Canada, the carbon tax for ethanol is equal to that for gasoline, which amounts to 6.67 cents per liter of fuel burned (B.C. Ministry of Finance, n.d.). Although the plant is being built in Ontario, carbon tax only exists for gasoline (not ethanol) in Ontario, which means that the baseline NPV does not include a carbon tax. However, it is possible that a carbon tax for ethanol equivalent to that for gasoline can be imposed in Ontario’s future, similar to B.C.’s carbon tax. This tax would then be 4.7 cents per liter of ethanol (Ontario Ministry of Finance, 2015). Figure 26 shows the sensitivity of the four factors mentioned. Ethanol Price

Energy Price

Plant Lifetime

Carbon Tax

$300.00

s n o i $200.00 l l i M $100.00 ] D A C [ V P N

$-100

-50

$(100.00) 0

50

100

$(200.00) $(300.00) $(400.00) $(500.00) $(600.00) $(700.00)

% Change

Fi gure 26 . A +20% sensitivity in ethanol price is an increase of 0.1 cents per liter. A +20% sensitivity in

feedstock price is an increase of 5 dollars per tonne. A +20% sensitivity in plant lifetime is an increase of 5 years (ranging from -40% to 40%). Carbon tax is a binary sensitivity with 0% representing NPV w/o carbon tax and 100% representing NPV w/ carbon tax.

The baseline NPV shown in Figure 26 is -$223 million for an ethanol price of 0.72 cents per liter, a feedstock cost of $71.5 per tonne of miscanthus, a 25 year lifetime and no carbon tax. Reiterating the point made in previous discussions, the primary reason why NPV is negative is due to the high feedstock cost, as the ethanol revenue generated 89


annually barely exceeds the annual feedstock operating cost ($90 million to $87 million). Figure 26 clearly shows that the only way for NPV to be positive is that the price of ethanol exceed about 1 dollar per liter. This represents 28 cents per liter or 1.06 dollars per gallon increase in the price. This is not only unlikely under current market conditions and variations but such a high price will turn off customers from buying the ethanol due to the lower energy content of ethanol blends when compared to pure gasoline. Therefore, realistically NPV can only become positive if the gasoline price goes back to normal levels from its current state, as it went through a downward spike around a year ago due to oversupply of crude oil. From the illustration in Figure 26, it can also be seen that a -100% sensitivity in feedstock price representing 46.5 dollars per tonne still cannot make the NPV positive. This shows how the plant can almost never be profitable, no matter how low the feedstock cost goes. It signifies the risk of introducing miscanthus to Ontario, with the associated risks of being an invasive species and requiring vast amounts of land to grow. The plant lifetime sensitivity in Figure 26 does not result in a positive NPV either. However, it is nonlinear and shows that a smaller plant lifetime can make the NPV significantly larger. Finally, the carbon tax sensitivity causes a further decrease in NPV as the tax takes away from the ethanol sales revenue. This is not important as the NPV is already negative and is not affected significantly by an added tax. The calculated incentive amount that was mentioned in Equation ### is exactly $7 million dollars. Though this is helpful in raising NPV, it is only a small amo unt and only useful when NPV is already above $0. The final NPV including the incentive over 2016 and 2017 is negative $217 million. This is still far below $0. Therefore, investment into this project is not recommended.

90


8.6 Equipment Sizing A list of each type of unit and how it was sized is shown below. The Aspen Plus simulation was used to size only a few units as the given sizes were unrealistic in most cases. Pretreatment – The belt dryer size was chosen based on relevant literature. It was 4-5 m

wide and 30 m long. It supported a maximum throughput of 500 tonnes/hr so it was enough to allow our process to run with no issues. The crusher size was based on the dryer, which meant that it would have to be around 4-5 m wide as well. As for the total volume of the crusher, the exact speed of crushing and the velocity of the belt dyer’s speed would have to be known in order to estimate. Additionally, we would have to make sure the two processes remain continuous when making adjustments to these specifications. Gasifier – The gasifier size was initially found using the Aspen Plus simulation, which

gave a 4.5 feet diameter as one of the parameters. Since this did not accurately fit the incoming feedstock size and amount, another method was used. In this method, the gas velocity out of the gasifier was decided as 5 m/s (typical) and the volumetric flow rate out of it was used to find the cross sectional area, giving a 20 m2 value. The diameter was then calculated as 5 meters. Usually, the height of the gasifier is longer than the diameter to account for the char and ash base, the incoming feedstock, the pyrolysis section and the fluidized bed. A reasonable height that was decided upon was 20 m high. Cyclone – It was assumed that an industrial cyclone with a comparable height to that of

the gasifier would have to be used since the gasifier’s outlet is fed into the cyclone. However, this did not have to be as high as the fluidized bed because a fluidized bed serves other functions except for moving/separating gas. Therefore, a 15 m height was chosen for the cyclone. The rest of the dimensions were proportional to this height. Heat exchangers – Heat exchangers were sized using the heat transfer area. To get a

detailed estimate, a diameter and length would have to be chosen for each pass of a shell and tube heat exchanger. The heat transfer area would then have to be divided by the surface area of a single pass to give the amount of passes required. Based on this, a rough 91


estimate of the total volume would be found. However, this detailed approach was not used as the pipe diameters and velocities inside and outside the HXs would have to be known or decided upon. E-101 = 620 m2 E-102 = 7543 m2 (this would have to be divided up into two or more HXs) E-103 = 529 m2 E-104 = 1224 m2 E-105 = 450 m2 E-106 = 463 m2 E-107 = 530 m2 Fermenter – A volume was picked based on literature (6000 m3), as explained previously. Vessels – Vessels (condensers) were sized using a length and diameter from the Aspen

Plus simulation. V-101 = 5.5 m length, 1.8 m diameter V-102 = 5.9 m length, 2 m diameter Pumps – Pump sizes had to be based on piping diameters. However, these were not set. In

addition, the volumes had to use the incoming volumetric flow amounts but the exact correlation was not known. It was therefore assumed that the pipe diameters were around 1 foot and that the pump sizes were proportional to this diameter. Distillation columns – Distillation columns were initially sized using Aspen Plus.

However, a 5 m diameter had to be set by us to account for transportation constraints on moving the columns, so the height of the columns were increased (using the same volume from Aspen Plus), which gave ~6 distillation columns for each single distillation column block, measuring 50-80 m each. These heights were slightly unreasonable but the values matched our smaller ethanol yield compared to theoretical values. T-101 = 405 m length (6 distillation columns / 67.5 m length each), 5 m diameter T-102 = 456 m length (6 distillation columns / 76 m length each), 5 m diameter Storage containers – Storage container volumes were based on a residence time of 3 days

and the given incoming flow rate. The storage containers were floating roof types. One

92


storage tank was actually the pre-fermentation mixer. Therefore, it had to be close to the fermenter size to maintain similar conditions. S-101 = 21,000 m3 S-102 = 6000 m3 Heat exchanger design Shell and tube countercurrent heat exchangers were simulated in Aspen plus in

order to determine the heat duty required to cool/heat a process stream to a specified outlet stream temperature using heat transfer coefficients from tabulated sources. Heat exchangers E-101, E-102 and E-103 were designed using equations and tabulated overall heat transfer coefficients from the textbook” chemical engineering design: principles practice and economic design” by Towler and Sinnott. The area of heat transfer was calculated using Equation 3. Meanwhile, the required heat transfer area for heat exchangers E-104, E-105, E-106 and E-107 were obtained using Aspen Plus Process Economic Analyzer.

=∆

Where,

[3]

Q = heat transferred per unit time [W] U = the overall heat transfer coefficient, [W/m2] A = heat transfer area [m2 ] ΔTm = the mean temperature difference [°C] The mean temperature difference was calculated using Equation 4;

Where ,

∆ = ∆

[4]

ΔTlm = log mean temperature difference [°C] and calculated using Equation 5; Ft = correction factor which depends on the heat exchanger geometry and obtained from (Towler and Sinnott, 2012).

Where,

 ) ∆ = ( ln ()( )    ( ) 93

[5]


T1 = hot fluid inlet temperature [°C] T2 = hot fluid outlet temperature [°C] t1 = cold fluid inlet temperature [°C] t2 = cold fluid outlet temperature [°C]

Heat Exchanger (E-101)

T1 = 850°C T2 = 550°C t1 = 37°C t2 = 134.976°C Q = 26,201,374.9 J/sec, obtained from Aspen Plus U = 70 W/m2 °C for a liquid water as cooling fluid in the shell and gas at 1 bar flowing inside the tubes.

976 °C)(550 °C37 °C) ∆ = (850 °C134. °C134. 9 76 °C) ln (850 (850 °C134.976 °C) =612.66°C

For a one shell and even multiple of tube passes, the correction factor, Ft, equals .9857. Therefore ΔTm is

∆ = .9857∙612.66 = 603.9°C  = ∆  26,201,374.9 = 70 W/m2 °C∙ 612.66°C 201,374.9 = 70 W26,°C∙ m2 612. 66°C = 619.8 

Rearranging, Equation 3 for A, the total heat transfer area required is then;

94


Specifications of heat exchangers E-101, E-103, E-104, E-105, E-106 and E-107 are summarized in Table 23.

Table 23. A summary of parameters used to design heat exchangers E-101, E-102 and E-103.

Equipment ID

E-101

E-102

E-103

Heat exchanger

Countercurrent Shell



type

and Tube

and Tube

and Tube

Q [J/sec]

26,201,374.9

12,000,822.4

5006737.3

70.0

70.0

800

T1 [°C]

850

550

73.2

T2 [°C]

550

187

37

t1[°C]

29.4

29.4

29.4

t2 [°C]

134.9

134.9

134.9

ΔTlm [°C]

612.6

23.0

14.3

Ft

.9857

0.8

0.825

ΔTm [°C]

603.9

18.4

11.8

A [m ]

619.8

9,299

529.3

U [W/m

∙°

C]

The heat transfer area for heat exchangers E-104, E-105, E-106, E-107 were obtained using Aspen Plus Process Economic Analyzer. The heat transfer area for heat exchangers E-104, E-105, E-106 and E-107 are 1224.177 m2, 450.2378 m2, 463.421 and 530.3198 m2 respectively.

9. Environmental Impact 9.1 LCA A cradle to gate LCA was decided in order to show the environmental impact from the various stages of the gasification-fermentation ethanol production process. This can be seen in Figure 27 below.

95


Figure 27. Cradle to gate LCA for the plant.

Starting from the production of the feedstock, the seeds and fertilizer are planted and the plant is grown over time in the presence of air, water and fertilizer before it is collected. During the growth of the feedstock, emissions are absorbed by the plant, however the equipment used on the farms in order to grow the feedstock are mostly powered with diesel from the refinery will have various emissions which are labeled as CO2, H2O, NOx, Sox (to air) on the LCA. Electricity from the grid will also contribute to the emissions on the farm. This results in the farm having a positive GHG value. Transportation emissions are due to using trucks to transport the feedstock from the farm to the plant and the trucks returning. For this project trucks of 44 tonne capacity are used as the vehicle since emission values are found in literature with 44 tonne trucks as basis (Bonitta and Whittaker, 2009). This contribute a large amount of emission since combustion of diesel results in high amounts of waste gas into the air. The diesel used by transportation and the farm comes from crude oil. The crude oil is taken out of the ground via drilling and is sent to the refinery to be refined and finished into usable fuel (gasoline, diesel, etc).

96


During the drilling phase, non-usable components of the oil is often combusted. Along with the production of the pipes, energy use from machines in drilling and refining, the process of diesel refining also contribute a positive GHG level. Another source of emission is from electricity generation. Since there are many forms of power plants, some contribute a much higher percentage of emissions than others. For example, gas power plants and coal power plants are older technologies which are banned from certain places due to their large environmental impact. Renewable energy sources such as wind and hydro have constantly been under research and development with the hope of replacing coal and gas one day. In terms of nuclear power plant, it is the cleanest reliable source of power today which resulted in nuclear power being the highest electricity production method to the grid (nuclear energy GHG emission: 16-55g CO2 equivalent/kWh, Gas energy GHG emissions: 700-1000g CO2 equivalent/kWh) (Fthenakis and Kim, 2007). Looking at the plant itself, there are emissions from the electricity use and fuel that may be consumed on site. The emissions from the plant can be found in the following section in Table 25. One reason the syngas fermentation method is said to emit less GHG is because the reactions themselves only produce two types of greenhouse gases in carbon dioxide and methane.

9.2 GHG Emissions When calculating the total GHG used in the syngas fermentation ethanol production process, the GHG emission from various steps in the LCA is calculated and added together. The LCA can be found in the previous section. The first portion of the LCA shows the farms that grows the feedstock used in the plant. For this project, Miscanthus is used as a representation of the feedstock however the process can be used with almost all types of biomass. Through research, it was found that for every kilogram of miscanthus grown, the total GHG emission is equivalent to 51 grams of CO2. (Maxime et al., 2013) Therefore since this process uses 122,728kg of miscanthus per hour, 6,259,128 g or 6,259kg of CO2 is released into the air per hour. Another source of GHG emission from the LCA is the transportation of feedstock from the farm to the plant. Assuming the trucks used for transportation are large diesel 97


trucks of 44 tonne capacity, meeting the usage demand of the plant would require up to 67 trucks every day to meet the required feedstock for the plant. From research articles, it was found these trucks emit equivalent of 120g CO2 per every kilometer traveled (Bonitta and Whittaker, 2009). Assuming the feedstock is available in a radius of 100km from the plant as well as the return trip for the trucks, the total GHG emissions from transportation would be equivalent to 1,608,000g or 1,608kg of CO2 per day for the trucks to go from the farms to the plant and then back. Since diesel is used for the truck transportation, the amount of energy used by each truck per trip is 356 MJ. Using correlation from research, 19.224 kg of CO2 equivalent is released per truck for each trip (Hsu, 2011). Taking account of the 67 trucks needed, the total CO2 equivalent comes out to be 1,288kg per day. The next source of emissions is from the generation of electricity. Since there has been insufficient research done on the farms supplying the feedstock thus far, only the electricity used for the production of ethanol in the plant will be included in the calculations. Since nuclear energy makes up of around two thirds of the or the electricity going to the grid, it is assumed the plant takes electricity purely from nuclear source since the supply of power through different methods (i.e. wind, hydro and solar) all vary depending on the time and weather of the day. Natural gas electricity accounts for a small portion of the power grid therefore it is neglected in this analysis. The high amount of nuclear energy is due to nuclear power plants are the most powerful and at the same time has the lowest GHG emissions from any of the non-renewable/non-natural methods of generating electricity (coal, natural gas). It is also reliable unlike methods such as wind, hydro and solar. Through research, it was found for every kilowatt/hour of electricity used, the total equivalent of CO2 was found to be between 16-55g (Fthenakis and Kim, 2007) . Assuming the Ontario nuclear power plants emit at the middle of the range, the total GHG was found to be 90,164.18g CO2 per day, or 90.16kg CO2 per day. Finally, in terms of the emissions from the plant itself, the environmental advantage of syngas fermentation over other types of ethanol production is showed by the difference in the GHG emissions. Looking at the results from the plant process, it was

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found only two types of greenhouse gas was produced in the process of fermentation (CO2 and CH4). Through simulations, the amount of CO2 in the products was found to be 6,188.09kg per hour and CH4 was produced at 354.57 kg per hour. The CO2 equivalent of CH4 produced was 12,055.4kg per hour was found using correlations found from research (Intergovernmental Panel on Climate Change ,2013). Due to inability to find related research, the emissions from diesel production from crude was neglected in the calculations of the emissions. Further research can be done in order to determine more detailed GHG emission and other methods of diesel production that may have lower emissions. Table 24 shows the CO2 equivalent of each source of GHG per unit capacity of the plant. The capacity of this plant is 303,031L of ethanol per day. Table 24. Cradle gate GHG emissions of ethanol produced.

CO2 Equivalent

Source

[kg/L-ethanol]

Feedstock Farm

0.496

Diesel Production

0.00531

for Transportation Transportation

0.00425

Electricity

0.000317

CO2 Produced by Plant

0.490

CH4 Produced by Plant

0.955

Total

1.945

As shown in Table 24, the highest amount of emissions from this process is the methane produced since it is assumed the methane would be flared into the atmosphere. However since methane is a usable resource, it could be separate out and used as a utility 99


or sold to customers if needed. It would increase the profitability of the plant greatly and will allow the plant to have lower emissions based on a cradle to gate analysis. In terms of the environmental benefits of bioethanol to gasoline, a LCA analysis for a pyrolysis gasoline production resulted in GHG emission of 22.34 kg CO2 Equivalent per liter of gasoline (Hsu, 2011) which is more than 10 times the amount found through a biomass syngas fermentation process shown in Table 24 above. Research showed potential of 88% decrease in GHG emission from using miscanthus as a biofuel feedstock (Huang et al., 2009) and this result agrees with that. This shows the improvement of a second generation ethanol production plant from a first generation plant in terms of emissions. GHG emission from burning bioethanol is much lower than burning gasoline. Using statistics from the United States Environmental Protection Agency, Motor gasoline had a total GHG emission of 2.329kg CO2 equivalent and the ethanol obtained total GHG emission of 1.52kg CO2 equivalent per liter. From those figures, just using ethanol instead of gasoline decreased the GHG by 35%. (EPA). Adding up the emissions from production and use, the total emission level from gasoline is 24.669 kg CO2 per liter and 3.465 kg CO2 per liter for bioethanol produced. This means an overall GHG reduction of around 85% by using bioethanol produced by Miscanthus syngas fermentation which comes close to values found from research. Just looking at those figures, replacing gasoline with ethanol will improve the environment greatly.

10. Process safety 10.1 Hazardous Materials Several chemicals are used as raw materials, produced in different sections of the plant or used as the ultimate products in the biochemical plant. The following chemicals listed in Table 25 pose safety hazards to both the equipment and personnel in the bioethanol plant.

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Table 25. List of hazardous chemicals used and produced in the biochemical plant along with a description of hazards they pose.

Material

Material Description

Incompatible materials

-Can cause asphyxiation by displacing oxygen -Exposure may cause headaches, drowsiness, vomiting etc. -LC50 of hydrogen is >15000 ppm/1h

-Can form explosive mixture with air and result to fire/explosion -Reacts violently with oxidants which leads to fire -Lower flammability limit : 4% -Upper flammability limit: 76% -Auto ignition temperature: 500 to 571°C

-Reactive with oxidizing materials

-Odourless -Colourless

-Highly toxic gas. -Causes severe asphyxiation and is toxic when inhaled -Over exposure may lead to loss of consciousness and muscle weakness -PEL of carbon monoxide is 50 ppm -LD50 of carbon monoxide is 1807 mg/kg

-Extremely flammable. -Must be isolated from sparks and source of heat -Lower flammable limit: 10.9% -Upper flammable limit of 79.2% -Auto ignition temperature of 607°C.

Reactive with oxidizing materials

-Odourless -Colourless

-May displace oxygen and cause suffocation at high concentration -PEL of carbon monoxide is 50000 ppm

-Odourless (< 5000 ppm) -Colourless

TLV of 1000 ppm

Gas

Carbon Monoxide

Dioxide

Methane

Combustibility

-Odourless -Colourless Hydrogen

Carbon

Flammability and

Toxicity

-Incompatible/Reactive with Magnesium, Titanium, Aluminum

-Ignites in presence of heat, sparks, open flames and hot surfaces -Decomposes to hazardous chemicals such as carbon dioxide and carbon monoxide. -Accidental release poses a serious fire or explosive hazard -Lower flammable limit of 1.8% an upper explosive limit of 8.4% -Auto ignition temperature of 287°C

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Reactive with oxidizing materials


Ammonia

-Colourless -Strong (pungent) -Unpleasant odour

-Causes extreme poisoning when inhaled -Causes serious eye damage and skin burns -Irritating and corrosive to the respiratory system -PEL of ammonia is 50 ppm -LC50 of 7338 ppm over an hour exposure

-Colorless -Acidic odour

-Causes severe burns to exposed skin -Extremely toxic to respiratory tract system -TLV of 2 ppm -LD50 of 4701 mg/kg.

-Greenishyellow colour -Pungent odour

-Nose and throat irritant -Chlorine is converted to hydrochloric acid in the respiratory system -Causes pneumonia and pulmonary enema. -Chlorine gas has a threshold limit value of 0.5 ppm

-Flammable gases will form explosive mixtures with chlorine

-Incompatible with several compounds including alcohols

-Colourless -Rotten eggs odour

-LC50 of 3124 ppm (1 Hour) -Extremely toxic and fatal when inhaled -TLV of 1 ppm -LC50 of 712 ppm over 1 hour -Inhalation or contact with ethanol can cause skin and eye irritation -PEL of ethanol is 1000 ppm -LD50 of ethanol is 3450(oral, mouse). -Corrosive -Causes severe irritation to the eyes and severe skin burns

-Extremely flammable in presence of sparks, heat source -Lower flammable limit: 4.3% -Upper flammable limit: 45% -Auto ignition temperature: 270°C

-Incompatible with oxidants such as oxygen difluoride

-Highly flammable and may form explosive mixtures with air -Lower flammable limit: of 3.3% -Upper flammable limit:19% -Auto ignition temperature of 423°C

-Reactive with oxidizing material

-PEL of acetic acid is 10 ppm

-Sulphur dust suspended in air is readily ignited by flame

-Readily ignites with air

Hydrogen Chloride

Chlorine Gas

Hydrogen Sulphide

Ethanol

Acetic acid

Sulphur

-Yellow color

-Lower flammable limit of 15.4% -Upper explosive limit of 25% -Auto ignition temperature 651°C. -Ignites in presence of oxidizing materials

-Reactive with oxidizing materials -Incompatible with halogens and acids

-Reactive with acetic anhydride, propylene oxide, sodium hydroxide

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LC50/LD50 = Lethal concentration/Lethal dose at which 50% of animal (tests) were killed TLV = Threshold limit value PEL = Permissible exposure limit

10.2 Process Hazards An advantage of producing ethanol through the hybrid process is the generally mild operating conditions. Fermentation requires atmospheric conditions and a temperature of 37°C which does not pose any extreme effects. On the other hand, gasification reaction operates at a high temperature of 850°C. This results to extreme temperatures that can melt tubes and pipes as well as cause ignition and start fire. In order to mitigate this risk, temperature and pressure control sensors and process control loops were pu t in place in order to counteract any sudden disturbances to temperature and pressure. Furthermore, the flow rate of the feed to the gasification is controlled in order to prevent runaway reactions. The preferred material of construction was high grade stainless steel. In addition to the gasifier reactor, safety control systems were integrated in to the distillation columns. Distillation columns pose the risk of high vapour flow rates which can lead to over pressurised distillation column and pressure vessels. This may cause explosions which can damage equipment and more importantly personnel. Temperature, pressure and flow rate control systems were put in place in order to mitigate these hazards. Furthermore, pressure relief systems were integrated into the pressure vessels which will release vapour to relieve high pressures. Duplicated temperature, pressure and flow rate sensors were also included in the design in order to safely and accurately measure any deviations. Automatic safety shutdown systems such as the use of solenoid valves which halts feed to equipment in case of unsafe conditions were also included in the plant design. Other equipment such as pumps, heat exchangers, valves are also integrated with control systems which counteracts any disturbances (which may pose safety hazards) to

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flow rates, temperature and/or pressure. A more detailed safety analysis of the d istillation columns is included in the HAZOP analysis in Appendix 3. Bypass streams were designed to handle increased flow rates, pipe leaks and prevent pipes from bursting caused by blockages. Storage vessels were also placed in key areas of the plant such as prior to the distillation column and fermentation vessel to not only increase reliability but to also store materials in case feed to eq uipment must be stopped to prevent safety hazards such as high level of material in the equipment. Other operating conditions include flushing equipment with an inert gas in order to prevent unwanted reactions caused by foreign materials. Finally, equipment are released of any gas by opening vent valves before commissioning and decommissioning.

11. Risk Assessment 11. 1 Technical One of the largest risks regarded with the syngas fermentation process is the bacteria used in fermentation. Since the bacteria is not yet commercialized or even used on a regular basis, the bacteria is completely unavailable for use on an industrial scale and is extremely expensive to purchase a culture from suppliers. As previously said in this report, one culture of these bacteria will cost over 70 dollars. In order to make this process feasible, growing the bacteria would be the most realistic option in order to provide the required amount of bacteria. Since the said bacteria do not have any history related to growing, various tests has to be done in order to see if it can be grown in large quantities for the use in the industry. If the bacteria cannot be grown, then this process is not a viable method of ethanol production. Another large risk technologically is the lack of knowledge and experience in the field of cellulosic bioethanol or even second generation bioethanol in general. As of now, Ontario does not have a full operating second generation bioethanol plant in operation (Decker, 2009). If this purposed plant were somehow to get approval, this would be the first plant ever. Because of the lack of industrial size plants, no information is known on the subjects of these plants in this region. For example, questions such as feedstock growth, temperature effects, demands and prices are all undetermined and since there are 104


no historical trends, establishing a knowledge base is extremely difficult which means a very extreme risk. Since there are no industrial scale syngas fermentation unit, creating a plant like this would require the designing of a unit without any prior knowledge which leads to very large uncertainties which may make it difficult to develop. There are also technological risks when growing the feedstock such as Miscanthus. From research, it may take up to five years in order to introduce the crop properly and to have desired growth in a new environment (Roy, 2014). Therefore if decided to start with the growth of Miscanthus, a solution that will speed up the growing process greatly would need to be find in order for this method to be feasible. The amount of feedstock needed for this plant is also very large and since there are always risks with crop growing in terms of climate, securing the required amount of feedstock every year can be difficult. This risk is not of large concern for this plant however since the syngas fermentation process is able to use almost every type of biomass available to produce ethanol. If not enough miscanthus can be grown or bought, other feedstock such as corn stover can be used in its place.

11.2 Societal Since this plant does not currently exist in Ontario, the effects of opening a new plant are not available to see. However, one of the risks of opening a new plant could be the potential impact on the workforce in the community. If the plant was opened in an area of low population, there would not be enough workers to supply the demand of the plant. Without workers, the plant would not produce anything, therefore the location of the plant is extremely important. Building a plant too close to large community may also have risks due to the response of the residents. Having a plant nearby may result in higher noise levels, higher emissions in the air and even potential pollution in the area. If these issues are not managed correctly, the residents could oppose the opening of the plant or the operation once built. If opposition were to occur, then it would lead to huge loss in the economics. This will be discussed further in the next section. Since the work force for each community is limited, adding further demand for workers in the community may take away jobs from other sectors in the community that 105


may result in lack of workers in those sectors. For example, workers from service sectors could see a job at the plant as a better choice and relocate themselves in the plant instead. By doing that, the workers in the service sector are greatly reduced, it may even cause a shortage of workers in service. This would lead to longer hours as well as more responsibilities than before. Unless the worker wages are increased, then the workers will be unsatisfied and important services such as caring for children and elders may take a hit in the total availability of workers. With the need t o grow the feedstock, more land will be required and in worst case scenarios it may require residents to relocate in order to free up the land needed which would cause opposition from residents. If relocation of residents is unavoidable, there would be additional costs on the start-up of the plant since there would need to be compensation for the relocated.

11.3 Economical Some of the economic risks can be seen in the sensitivity analysis. Since the most important part of economics is the profit, the goal must always be to maximize the profit of the plant. Looking at the sensitivity analysis, the largest factor that will affect profitability is the ethanol sales price. Since the price of ethanol fluctuates and introducing more ethanol into the market will likely to decrease the price of ethanol at the beginning, the operation is at risk of losing profits especially at the start of operation. The second most significant factor explored on the sensitivity analysis is the Energy price. In this case, the energy price is directly related to the electricity cost of running the plant. Since profit is relatively sensitive to the energy price, an y increase in the energy price will result in loss of profits. Looking at the recent trends of energy prices in Ontario, the price has been increasing steadily over the past years (stats Canada) therefore assuming the energy prices will continue in that trend, there is a risk of profit much lower than anticipated once the plant is in production. Another concern economically would be the demand of ethanol. In estimations, the market for ethanol is assumed to be open where there are no barriers to entry. In this market, any cowpony that chooses to produce ethanol will be able to enter the market without any restrictions which means the market could potentially become 106


oversaturated with supply and not enough demand to keep the price at an acceptable rate. A market such as this has the potential to drive the market price below the profitable level for every supplier in the market which will lead stop in operation for many plants. While the government can demand a production limit to every plant in order to prevent the market from deteriorating, larger plants will require a certain level of production in order to stay profitable and may simply decide to not operate under the constraints while smaller companies will not be able to match the total demand by themselves. This would then raise the price of ethanol and also raise the unemployment in the area due to plant shutdowns (Vazirani, 2007). Since the gasoline industry is so large and has large amounts of influence in the world, it will not be easy to even secure a market in order to become profitable in ethanol production. Even now, there are many uses of gasoline that cannot be replaced with ethanol and at best, most of the machines require a blend of ethanol and gasoline in order to function. In order for the production process to be profitable and ensure sufficient demand, further ethanol uses will need to be developed and ethanol will need to be integrated as a viable source of energy along with gasoline. The government will likely have to intervene in order to establish the base for ethanol to become successful economically.

107


Appendices Appendix 1 – Various Lists Relating to Process List of Materials Adsorbent (Zinc Oxide [ZnO]) Bacteria (Clostridium Ijungdahlii) Corn Steep Liquor (Fermentation Medium) Feedstock (i.e. Miscanthus, Switchgrass, Corn Stover etc.) Natural Gas Nutrients Oxygen Water        

List of Equipment Adsorption Vessels Conveyer Belts Cyclone Distillation Columns Fluidized Bed Gasifier Gravity Chute Heat Exchangers Pumps Screw Belts Sensors Stirred-Tank Bioreactor Storage Tank (Distillation Column Feed) Storage Tank (Feedstock) Storage Tank (Fermentation Medium) Valves (i.e. Manual, Automatic) Wet Scrubber (Spray Tower)                

List of Symbols Q – Heat transferred/time [W] U – Overall heat transfer coefficient [W/m2]  

108




A – Heat transfer area [m2] ΔTm – Mean temperature difference [⁰C]



ΔTlm – Log-mean temperature difference [⁰C]



    

Ft – Correction factor T1 – Hot fluid inlet temperature T2 – Hot fluid outlet temperature t1 – Cold fluid inlet temperature t2 – Cold fluid outlet temperature

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Appendix 2- Detailed Equipment List Equipment List Displayed Text

Type

Specific

B-101

Conveyer Belt

B-102

Conveyer Belt

B-103

Belt

Conveyer Belt

B-104

Screw Belt

B-105

Screw Belt

C-101

Chute

E-101

Rubber

Electricity

Insulation tip

Heat Exchanger

Cooling recycle from T-101 before S-101 Reboiler for T-101

None Cooling Water Cooling Water

Cooling of syngas during gas cleaning

E-103 E-105

Utility

Cooling syngas from R-101

E-102 E-104

Gravity Chute

Extra

Cooling Water Shell & Tube

Cooling Water Low Pressure Steam

E-106

Condenser for T-101

Cooling Water

E-107

Reboiler for T-102

Low Pressure Steam

E-108

Condenser for T-102

GC-101

Cyclone

None

None

GC-102 A

Adsorption Column

Adsorbent (ZnO)

Oxygen

GC-102 B

Gas Cleaning

Cooling Water

Adsorption Column

Adsorbent (ZnO)

Oxygen

GC-103

Wet Scrubber

Spray Nozzle

Cooling Water

P-102 A

Pump broth from R-102 to T-101

None

Electricity

Hammer Mill

None

Electricity

Spray Dryer

Conveyer Belt

Steam

Gasifier

Fluidized Bed (Quartz-Sand)

Oxygen/Natural Gas

Fermenter

Bacteria (Clostridium ljungdahlii)

Electricity/Water

P-101 A

Pump new medium i nto R-102

P-101 B P-102 B

Pump broth from R-102 to T-101

P-103

Pump from S-102 to T-101

P-104 A P-104 B

Pump

P-105 A

Pump liquid from V-101

P-105 B P-106 A

Pump bottoms from T-102

P-106 B P-107 A

Pump liquid from V-102

P-107 B PT-101 PT-102 R-101 R-102 S-101 S-102 T-101 T-102 V-101 V-102

Pump bottoms from T-101

Pre-Treatment Reactor Storage Tower Vessel

Medium Storage

None

None

Column feed storage

None

None

Distillation Column

Trays

None

Distillation Column

Trays

None

Condenser vessel for T-101

None

None

Condenser vessel for T-102

None

None

Figure A2.1. Detailed equipment list.

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Appendix 3 - HAZOP Study Node – T-101 – Distillation Column Parameters – Flow, Temperature, Pressure Guide Words – Low, High, No (if applicable) Parameter: Flow Guide Word

Deviation

Low liquid column flow Low

Low vapour column flow

High liquid column flow

Cause

Consequence

- Fouling in the pipe - Faulty level control loop from reactor, causing feed flow to decrease - Reboiler duty too high - Condenser duty too low - Pump breakdown on reflux loop

- Damage the pumps surrounding the column - Poor separation quality - Not meeting production requirements

- Column level control loop using distillate flow - Reflux and reboiler flow controls can also help raise liquid flow - Flow indicators placed at inlets and outlets of column

- Reflux flow too high due to temperature control loop, or condenser vessel level control loop - Reboiler duty too low, or boiler feed water flow/temperature too low - Faulty level control on fermenter, feeding excess liquid from reactor - Reflux too high from temperature control or condenser level control

- Can lead to weeping and even dumping of the trays - Reduces separation efficiency

- Control loop within reboiler controlling utility flow in ratio to the feed flow - Reboiler flow also in ratio control with bottoms flow to minimum vapour flow met

- Reduce separation quality - Can lead to flooding of condenser vessel - Can damage pumps

- Reboiler flow too high from column level control - Reflux too low - Condenser duty too low or cooling water

- Can knock trays out of place - Can cause flooding within the column

- Column level alarms - Reflux flow control loop - If feed flow too much for column to handle, send feed to storage tanks (done using control loop) - Level alarms placed on the column - Reflux control loop as preliminary control to prevent rising vapour flow

High

High vapour column flow

Action


flow too low

No liquid column flow

No

No vapour column flow

- Leads to entrainment - Tray efficiency reduced - Increase pressure within column - Breakdown of pumps surrounding column - Column will not operate, no separation

- Broken valve/piping - Reactor failure/shutdown - Fermenter flash vessel rupture

- Reboiler broken or fouled - Broken pipe in reboiler loop - Boiler water issues (temperature too low, or control causes low flow) - Reboiler ratio error causing no flow to go to reboiler

- Poor separation, reduced product quality, poor column performance - Can lead to weeping or dumping of trays

- Feed sent to storage tanks until distillation column is back and running properly - Parallel pumps used around column with alternate motors in case one pump runs dry or gets damaged - Control loop placed on the reboiler loop controlling utility flow in ratio to the feed flow entering the column

Parameter: Temperature Guide Word

Deviation

Low

Low column temperature

High

High column temperature

Cause - High liquid flow through the column - Low vapour flow through the column - Feed flow is too large - Reboiler duty too low - Reflux flow too high - High vapour flow within the column, or low liquid flow - Vapour condenser not working properly

112

Consequence - Poor column operation and product purity directly influenced by column temperature

- Can lead to unsafe conditions and poor column operation causing product purity to suffer

Action - Temperature control loop added using reflux flow rate to control column temperature and distillate flow rate - Two alternate temperature sensors were used sending the average value to the controller - Temperature indicators were placed at the top and bottom of the column, as well as near the feed tray.


Parameter: Pressure Guide Word

Low

High

Deviation

Cause

Consequence

- Reboiler duty is too low - Feed flow is too low - Reflux flow too high

- Can lead to weeping or dumping of trays - Column needs to be restarted if dumping occurs - Decrease in separation efficiency

- If vapour flow is too high into the column from reboiler loop - If condenser duty is too low - Outlet valve is broken or malfunctioning - Low/no liquid flow into column

- Can cause entrainment, reduce tray efficiency - Can eventually lead to flooding in the column - Significant decrease in separation efficiency

Low column pressure

High column pressure

113

Action

- Pressure within column is controlled via utility flow into the condenser - Two alternate pressure sensors used to increase reliability, where the average reading between the two sensors is sent to the controller - Low pressure alarm placed near top of column - Pressure relief is achieved via a safety relief valve near the top of the column in case of pressure buildup - High pressure alarm placed near the top of the column - Pressure indicators placed throughout the column

Figure. A3.1 HAZOP changes reflected on a P&ID.

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Bio-Ethanol Productions Plant in Ontario

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