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 – Chemical Plant Design and Simulation

4W4 2015 – Syngas Fermentation

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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 Canadian General Standards Board. The production of Ethanol 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 both the thermochemical and biochemical 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, and high resistance to contaminants. The feedstock of interest is first

crushed and dried before it is sent to a fluidized bed gasifier to 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 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 roughly $110 million per year to operate

and maintain. Though due to high feedstock costs incurred from growing 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, comparable to ~20 kg for crude oil processes. Though it 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 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

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 – 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 – 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 13

Figure 2. Length of Growing Season in Ontario 21

Figure 3. Block flow diagram of syngas fermentation process 22

Figure 4. Typical fluidized bed gasifier configuration 25

Figure 5. Typical biomass feeding system for fluidized bed gasifier 26

Figure 6. Typical wet scrubber configuration 28

Figure 7. Wood-Ljungdahlii biochemical pathway 30

Figure 8. Typical stirred-tank bioreactor configuration 31

Figure 9. Ratio control loop design for steam to feed ratio 59

Figure 10. Ratio control design for Feed to reboiler utility ratio 61

Figure 11. Ratio control structure between distillate and reflux 62

Figure 12. Pressure control inside the gasifier unit 64

Figure 13. Pressure control loop design for the distillation columns 65

Figure 14. Level control structure for the fermenter 66

Figure 15. Level control for reflux drum 67

Figure 16. Cascaded temperature control design around the condenser E-104 69

Figure 17. pH control loop structure for the fermenter. 71

Figure 18. Ratio control structure between purge stream and recycle stream 72

Figure 19. pH control design for unit S-101 73

Figure 20. Pre-treatment section 75

Figure 21. Gasification section 75

Figure 22. Gas cleaning section of the P&ID 76

Figure 23. Summary of Fermentation section of P&ID 77

Figure 24. Summary of Distillation section of the process 77

Figure 25. NPV Analysis 88

Figure 26. Sensitivity Analysis 90


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Table of Tables Table 1. Economic analysis of a cellulosic ethanol plant using the biochemical process 18

Table 2. Economic analysis of a cellulosic ethanol plant using the thermochemical process 20

Table 3. Process side material inflows and outflows of the bioethanol plant 41

Table 4. Utilities side overall material inflows of the bioethanol plant 42

Table 5. Utilities side overall material outflows of the bioethanol plant 43

Table 6. Overall energy balance of the process streams 44

Table 7. Overall energy inflow of the utility streams 45

Table 8. Overall energy inflow of the utility streams 46

Table 9. Stream table of process streams entering and exiting the pre-treatment section 46

Table 10. Material and energy inflow and outflow to the equipment of the pre-treatment 47

Table 11. Stream table of process streams entering and exiting gasification section 48

Table 12. Material and energy inflow and outflow to the equipment of the gasification section 49

Table 13. Stream table of process streams entering and exiting the gas cleaning section 50

Table 14. Material and energy inflow and outflow to the equipment of the gasification section 51

Table 15. Stream table of process streams entering and exiting the fermentation section 52

Table 16. Material and energy inflow and outflow to the equipment of the fermentation section 53

Table 17. Stream table of process streams entering and exiting the separation section 54

Table 18. Material and energy inflow and outflow to the equipment of the separation section 55

Table 19. Stream table of utilities streams entering and exiting the pre-treatment section 56

Table 20. Utilities side material and energy inflow and outflow to equipment of the gasifier 57

Table 21. Sum of capital costs for each type of unit and total capital cost. 81

Table 22. All operating costs for the syngas fermentation plant. 84

Table 23. Cradle gate GHG emissions of ethanol produced. 99

Table 24. List of hazardous chemicals used and produced in the biochemical plant 101

Table of Reacions Reactions 1-3. Reactions taking place within gasifier 25

Reactions 4-10. Primary reactions that occur within fermenter, dependant on H2/CO content 30

Reaction 11. Regeneration reaction of adsorbent bed. 34

Table of Equations Equation 1.Price of E100 86

Equation 2. Incentive calculation 88

Equation 3. Heat transfer area 91

Equation 4. Mean Temperature Difference 91

Equation 5. Log-Mean Temperature Difference 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


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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.


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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 0.36-

0.39 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.


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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 CO2 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 pre-

treatment 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


19

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).


20

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


21

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).


22

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 bio-

waste (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.


23

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.


24

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.


25

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.

𝐶 + 𝑂2

→ 𝐶𝑂2 [1]

𝐶 + 𝐻2𝑂

→ 𝐻2 + 𝐶𝑂 [2]

𝐶𝑂 + 𝐻2𝑂

→ 𝐶𝑂2 + 𝐻2 [3]

Reactions 1-3. Reactions taking place within gasifier. 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

drawn from the gasifier at an approximate temperature of 850⁰C, as is the produced ash

from the bottom of the unit.


26

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


27

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

~1000⁰C (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

the syngas must be cooled to an approximate temperature of 600⁰C 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

reach an approximate temperature of 45⁰C, all the while recovering significant amounts


28

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).

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

37⁰C 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.


29

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

be held as close to 37⁰C 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. A simplified reaction pathway is

shown below in Figure 7.


30

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).

6𝐶𝑂 + 3𝐻2𝑂

→ 𝐶𝐻3𝐶𝐻2𝑂𝐻 + 4𝐶𝑂2 [4]

4𝐶𝑂 + 2𝐻2𝑂

→ 𝐶𝐻3𝐶𝑂𝑂𝐻 + 2𝐶𝑂2 [5]

3𝐶𝑂 + 3𝐻2

→ 𝐶𝐻3𝐶𝐻2𝑂𝐻 + 𝐶𝑂2 [6]

2𝐶𝑂 + 2𝐻2

→ 𝐶𝐻3𝐶𝑂𝑂𝐻 [7]

2𝐶𝑂 + 4𝐻2

→ 𝐶𝐻3𝐶𝐻2𝑂𝐻 + 𝐻2𝑂 [8]

2𝐶𝑂2 + 6𝐻2

→ 𝐶𝐻3𝐶𝐻2𝑂𝐻 + 3𝐻2𝑂 [9]

2𝐶𝑂2 + 4𝐻2

→ 𝐶𝐻3𝐶𝑂𝑂𝐻 + 2𝐻2𝑂 [10]

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


31

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).


32

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


33

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


34

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:

𝑍𝑛𝑆 + 3

2𝑂2 → 𝑍𝑛𝑂 + 𝑆𝑂2 [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


35

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


36

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


37

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


38

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.


39

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


40

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 vapour), 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.

6.1.1 Process Side

Table 3. Process side material inflows and outflows of the bioethanol plant

PROCESS SIDE – OVERALL MATERIAL BALANCE

INFLOW OUTFLOW

Component Raw

Mat

Steam Fresh

Broth

O2-

H2O

NH3

Removal

Waste WWT DC-2

Bottom

Ethanol Evaporated

Water

Cyclone

underflow

H2S

removed

Total 122,728 93,992 31,040 70,846 735 127,650 952 38,406 8,819 17 218 119

Total mass

flow 247,760 247,762

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


42

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 DRYER-

STEAM CW1 CW2 CW3 CW4

DIST1-CW

DIST2-

CW

BOTTOMS1-

LPS BOTTOMS2-LPS

Total 69,945 153,766 89,667 205,641 6,469,180 1,490,110 827,897 60,645 27,262

Total mass

flow 9,520,617

Table 5. Utilities side overall material outflows of the bioethanol plant

UTILITIES SIDE- OVERALL MATERIAL OUTFLOWS

Component

DRYER-

STEAM

EXIT

LPS1 LPS2 CWR3 CWR4 DIST1-CW

RETURN

DIST2-CW

RETURN

BOTTOMS1

-LPS

RETURN

BOTTOMS2-LPS

RETURN

Total 69,945 153,766 89,667 205,641 6,469,180 1,490,110 827,897 60,645 27,262

Total mass flow 9,520,617


43

6.2 Overall Energy Balance

Table 6 displays the overall energy 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 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 OUTFLOW

Energy Raw Mat Steam Fresh

Broth O2-H2O

NH3

Removal Waste WWT

DC-2

Bottom Ethanol

Evaporated

Water

Cyclone

underflow

H2S

removed

Total Fluid

Enthalpy Flow

-356,727 -1,099,100 -495,460 -16,262 -1,963 -61,705

-15,014 -575,230 -52,071 -10,382

-217,910 -1.7

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 section, fermentation unit and the


44

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, except temperature which is in units of °C and pressure in units of atm

UTILITIES SIDE: ENERGY INFLOW

Energy CW1 CW2 CW3 CW4 DRYER-STEAM MILLING-

ELEC

Total Fluid Enthalpy

Flow -2,451,400 -1,427,900 -3,278,400 -103,130,000 73,485

Electricity .0025

Energy DIST1-CW DIST2-CW DIST1-LPS DIST2-LPS P-101

ELECTRICITY

P-102

ELECTRICITY

Fluid Enthalpy Flow -23,737,000 -13,188,000 -937,570 -421,480

Electricity 8x10-4

1.3x10-3

Energy P-103

ELECTRICITY

P-104

ELECTRICITY

P-105

ELECTRICITY

P-106

ELECTRICITY

P-107

ELECTRICITY

Fluid Enthalpy Flow 1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3


45

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

Energy LPS1 LPS2 CWR3 CWR4

DRYER-

STEAM

EXIT

Fluid Enthalpy Flow -2,356,900 -1,384,700 -3,260,300 -102,570,000 69,302

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

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


46

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

Component 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

Total Flow 122,728 17 122,711 122,711 122,711

Total Enthalpy

Flow -356,727 -10,382 -350,498 -350,498 -350,498


47

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] PT-102 [HAMMER MILL]

IN OUT 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


48

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

Component 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

Total Flow 93,992 122,711 145,859 70846

Total Enthalpy Flow -1,099,100 350,498 -342,170 -16,262


49

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

Total Flow 93,992 122,711 145,859 70,846

Total Enthalpy Flow -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.


50

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

Component SYNGAS CYCLONE

OVERFLOW

CYCLONE

UNDERFLOW

PRE-

COOLED

SYNGAS

CLEAN

SYNGAS

H2S

REMOVED

COOLED

SYNGAS

FERMENTER

FEED

NH3

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

Total Enthalpy

Flow -342,170 -560,080 217,910 -654,220 -436,300 -1.7 -587,620 -585,660 -1,963


51

Table 14. Material and energy inflow and outflow to the equipment of the gasification section of the bioethanol plant.

GC-101 [CYCLONE] E-101

[Heat exchanger]

GC-102 A/B

[Adsorption column]

E-102

[Heat exchanger]

GC-103

[Ammonia scrubber]

Component IN OUT IN OUT IN OUT OUT IN OUT IN OUT OUT

H2 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

Total Flow 145,859 145,641 218 145,641 145,641 145,641 119 145,522 145,522 145,522 145,522 144,787 0

Total Enthalpy

Flow -336,600 -560,080 -217,910 -560,080 -645,220 -645,220 -1.7 77,605

-

436,300

-

587,620

-

587,620

-

585,660 -1,963

Table 15 and Table 16 shows the stream table and equipment table for the fermentation section of the bioethanol plant.



52

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

Component FERMENTER

FEED

FRESH

BROTH

RECYLED

BROTH

MIXED

BROTH

COOLED

MIXED

BROTH

PUMP

MIXED

BROTH

FERMENTER

PRODUCT

DC-1

FEED WASTE

H2 10,112 0 0 0 0 0 0 0 315

CH4 355 0 0 0 0 0 0 0 355

CO 121,799 0 0 0 0 0 0 0 0

CO2 11,709 0 0 0 0 0 0 0 6188

H2O 0 31,040 56,531 87,571 87,571 87,571 93,274 93,274 308

NH3 0 0 0 0 0 0 0 0 0

HCL 812 0 0 0 0 0 0 0 812

H2S 0 0 0 0 0 0 0 0

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 10,398 119,579

Total Flow 144,787 31,040 65,003 96,043 96,043 96,043 113,181 113,181 127,650

Total Enthalpy

Flow -585,660 -495,460 -942,330 -1,437,800 -1,455,800 -1,455,800 -1,614,500 -1,614,500

-61,705


53

Table 16. Material and energy inflow and outflow to the equipment of the fermentation section of the bioethanol plant.

R-102 [Fermentation vessel] E-103 [Heat exchanger] P-101 A/B [PUMP] P-102 A/B [PUMP]

Component IN OUT IN OUT IN 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

H2S 0 0 0 0 0 0 0 0 0

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.



54

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

Component DC-1

FEED

DC-1

BOTTOM

DC-1

HEATED

BOTTOM

DC-1

Boilup

RECYCLED

BROTH DC-1 TOP

DC-1 TOP

CONDENS

DC-1

VESSEL

EXIT

DC-1

REFLUX DC-1 DISTILLATE

H2O 93,274 113,062 113,062 56,531 56,531 40,417 40,417 40,417 3674.3 36,743

Ethanol 9,509 `12 `12 6 6 10,452 10,452 10,452 950.2 9,502

Acetic Acid 10,398 16,934 16,934 8,467 8,467 2,125 2,125 2,125 193.2 1,932

Total Flow 113,181 130,008 130,008 65,003 65,003 52,994 52,994 52,994 4817.7 48,177

Total Enthalpy

Flow -1,614,500 -942,330 -942,330 -64,372 -643,720

Component DC-2

BOTTOM

DC-2

HEATED

BOTTOM

DC-2

Boilup

DC-2

BOTTOM DC-2 TOP

DC-2 TOP

CONDENS

DC-2

VESSEL

EXIT

DC-2

REFLUX

DC-2

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

Total Enthalpy

Flow -281,863 -402,661 -52,412 -66,673 -52,071

-

15,014


55

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

T-101 E-104 P-104 E-105 P-105

Component IN OUT IN OUT IN OUT IN 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 8,467 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 52,994 52,994 52,994

Total Enthalpy Flow -1,614,500 -942,330 -643,720

T-102 E-106 P-106 E-107 P-107

Component IN OUT IN OUT IN OUT IN 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 1,161 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

Total Enthalpy Flow -643,720 -575,230 -66,673

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. Component and total flows are per mass basis with units of kg/hr while total enthalpy flows are in

units of MJ/hr.


56

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

Component DRYER-

STEAM

DRYER-

STEAM EXIT

CW1 LPS1 CW2 LPS2 CW3 CW3 RETURN

H20 69,945 69,945 153,766 153,766 89,667 89,667 205,641 205,641

Total

Enthalpy

Flow

-73,485 -69,302 -2,451,400 -2,356,900 -1,427,900 -1,384,700 -3,278,400 -3,260,300

Component CW4 CWR4 DIST1-CW DIST1-CW

RETURN

DIST2-CW DIST2-CW

RETURN

DIST1-LPS DIST1-LPS

RETURN

DIST2-

LPS

H20 6,469,180 6,469,180 1,490,110 1,490,110 827,897 827,897 60,645 60,645 27,262

Total

Enthalpy

Flow

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

Component DIST2-

LPS

RETURN

P-101

ELECTRICITY

P-102

ELECTRICITY

P-104

ELECTRICITY

P-105

ELECTRICITY

P-106

ELECTRICITY

P-107

ELECTRICTY

MILLING

ELECTRICITY

H20 827,897

Total

Enthalpy

Flow -480,871

Electricity 8x10-4

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

.0253


57

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

Component PT-101 PT-102 E-101 E-102 E-103

H20 69,945 69,945 153,766 153,766 89,666 89,666 205,641 205,641

Total

Enthalpy

Flow

73,485 69,302 -2,451,400 -2,356,900 -1,427,900 -1,384,700 -3,278,400 -3,260,300

Electricity .0253

Component E-104 E-105 E-106 E-107 P-101

H20 56,531 56,531 1,490,110 1,490,110 27,262 27,262 827,897 827,897

Total

Enthalpy

Flow

-937,570 -1,069,866 -23,737,000 -23,633,731 -421,480 -480,871 -13,188,000 -13,130,424

Electricity 8x10-4

Component P-102 P-103 P-104 P-105 P-106 P-107

Electricity 1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

1.3x10-3

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


59

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.


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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


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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 temperature 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


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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)


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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.


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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


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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


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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


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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.


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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


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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


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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


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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


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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%.


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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


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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


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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 negligible.

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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Table 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 $ 2,374,000

TOTAL $ 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


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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


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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


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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.

Table 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 $ 36,000


84

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)

$ 1,870,000

TOTAL

$ 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


85

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


86

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).

𝑃𝑟𝑖𝑐𝑒 𝑜𝑓 𝐸100 = (

$2.21

𝑔𝑎𝑙𝑙𝑜𝑛 𝑜𝑓 𝐸85−0.15·

$2.30

𝑔𝑎𝑙𝑙𝑜𝑛 𝑜𝑓 𝑔𝑎𝑠𝑜𝑙𝑖𝑛𝑒

0.85) ·

1.25 𝐶𝐴𝐷

1 𝑈𝑆𝐷=

$2.74 𝐶𝐴𝐷

𝑔𝑎𝑙𝑙𝑜𝑛 (1)


87

Period All Income All eligible expenses All ineligible expenses Book value Depreciation Taxable income Tax paid Net cash flow in period (TVM) NPV

0 -$ -$ (87,129,306)$ 87,129,306$ 13,069,396$ (13,069,396)$ (3,267,349)$ (83,861,957)$ (83,861,957)$

1 90,000,207$ (112,713,359)$ -$ 74,059,910$ 22,217,973$ (44,931,125)$ (11,232,781)$ (10,436,700)$ (94,298,658)$

2 90,000,207$ (112,713,359)$ -$ 51,841,937$ 15,552,581$ (38,265,733)$ (9,566,433)$ (10,865,057)$ (105,163,714)$

3 90,000,207$ (112,713,359)$ -$ 36,289,356$ 10,886,807$ (33,599,958)$ (8,399,990)$ (10,753,690)$ (115,917,405)$

4 90,000,207$ (112,713,359)$ -$ 25,402,549$ 7,620,765$ (30,333,916)$ (7,583,479)$ (10,333,770)$ (126,251,175)$

5 90,000,207$ (112,713,359)$ -$ 17,781,784$ 5,334,535$ (28,047,687)$ (7,011,922)$ (9,749,228)$ (136,000,403)$

6 90,000,207$ (112,713,359)$ -$ 12,447,249$ 3,734,175$ (26,447,326)$ (6,611,832)$ (9,088,775)$ (145,089,178)$

7 90,000,207$ (112,713,359)$ -$ 8,713,074$ 2,613,922$ (25,327,074)$ (6,331,768)$ (8,406,240)$ (153,495,418)$

8 90,000,207$ (112,713,359)$ -$ 6,099,152$ 1,829,746$ (24,542,897)$ (6,135,724)$ (7,733,492)$ (161,228,910)$

9 90,000,207$ (112,713,359)$ -$ 4,269,406$ 1,280,822$ (23,993,974)$ (5,998,493)$ (7,088,647)$ (168,317,557)$

10 90,000,207$ (112,713,359)$ -$ 2,988,585$ 896,575$ (23,609,727)$ (5,902,432)$ (6,481,260)$ (174,798,817)$

11 90,000,207$ (112,713,359)$ -$ 2,092,009$ 627,603$ (23,340,754)$ (5,835,189)$ (5,915,623)$ (180,714,440)$

12 90,000,207$ (112,713,359)$ -$ 1,464,406$ 439,322$ (23,152,474)$ (5,788,118)$ (5,392,837)$ (186,107,278)$

13 90,000,207$ (112,713,359)$ -$ 1,025,084$ 307,525$ (23,020,677)$ (5,755,169)$ (4,912,123)$ (191,019,401)$

14 90,000,207$ (112,713,359)$ -$ 717,559$ 215,268$ (22,928,419)$ (5,732,105)$ (4,471,640)$ (195,491,041)$

15 90,000,207$ (112,713,359)$ -$ 502,291$ 150,687$ (22,863,839)$ (5,715,960)$ (4,068,993)$ (199,560,034)$

16 90,000,207$ (112,713,359)$ -$ 351,604$ 105,481$ (22,818,633)$ (5,704,658)$ (3,701,544)$ (203,261,578)$

17 90,000,207$ (112,713,359)$ -$ 246,123$ 73,837$ (22,786,988)$ (5,696,747)$ (3,366,605)$ (206,628,183)$

18 90,000,207$ (112,713,359)$ -$ 172,286$ 51,686$ (22,764,837)$ (5,691,209)$ (3,061,546)$ (209,689,729)$

19 90,000,207$ (112,713,359)$ -$ 120,600$ 36,180$ (22,749,332)$ (5,687,333)$ (2,783,857)$ (212,473,586)$

20 90,000,207$ (112,713,359)$ -$ 84,420$ 25,326$ (22,738,478)$ (5,684,619)$ (2,531,183)$ (215,004,769)$

21 90,000,207$ (112,713,359)$ -$ 59,094$ 17,728$ (22,730,880)$ (5,682,720)$ (2,301,332)$ (217,306,101)$

22 90,000,207$ (112,713,359)$ -$ 41,366$ 12,410$ (22,725,561)$ (5,681,390)$ (2,092,283)$ (219,398,384)$

23 90,000,207$ (112,713,359)$ -$ 28,956$ 8,687$ (22,721,838)$ (5,680,460)$ (1,902,180)$ (221,300,564)$

24 90,000,207$ (112,713,359)$ -$ 20,269$ 6,081$ (22,719,232)$ (5,679,808)$ (1,729,320)$ (223,029,884)$

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

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


88

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 ($

L) [2]

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


89

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.

Figure 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

$(700.00)

$(600.00)

$(500.00)

$(400.00)

$(300.00)

$(200.00)

$(100.00)

$-

$100.00

$200.00

$300.00

-100 -50 0 50 100

NP

V [

CA

D]

Mill

ion

s

% Change

Ethanol Price Energy Price Plant Lifetime Carbon Tax


90

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 amount 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.


91

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


92

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


93

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.

𝑄 = 𝑈𝐴∆𝑇𝑚 [3]

Where,

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;

∆𝑇𝑚 = 𝐹𝑡∆𝑇𝑙𝑚 [4]

Where ,

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).

∆𝑇𝑙𝑚 =(𝑇1 − 𝑡2) − (𝑇2 − 𝑡1)

ln(𝑇1 − 𝑡2)(𝑇2 − 𝑡1)

[5]

Where,


94

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.

∆𝑇𝑙𝑚 =(850 °C − 134.976 °C) − (550 °C − 37 °C)

ln(850 °C − 134.976 °C)(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

∆𝑇𝑚 = .9857 ∙ 612.66

= 603.9°C

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

𝐴 =𝑄

𝑈∆𝑇𝑚

=26,201,374.9

70 W/m2 °C ∙ 612.66°C

=26,201,374.9

70W

m2 °C ∙ 612.66°C

= 619.8 𝑚2


95

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

type

Countercurrent Shell

and Tube


and Tube


and Tube

Q [J/sec] 26,201,374.9 12,000,822.4 5006737.3

U [W/m2∙ °C] 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 [m2] 619.8 9,299 529.3

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 m

2, 463.421 and 530.3198 m

2

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.


96

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).


97

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


98

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


99

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.

Source CO2 Equivalent

[kg/L-ethanol]

Feedstock Farm 0.496

Diesel Production

for Transportation 0.00531

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


100

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.


101

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

of hazards they pose.

Material Material

Description Toxicity

Flammability and

Combustibility Incompatible materials

Hydrogen

Gas

-Odourless

-Colourless

-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

Carbon

Monoxide

-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

Carbon

Dioxide

-Odourless

-Colourless

-May displace oxygen and

cause suffocation at high

concentration

-PEL of carbon monoxide is

50000 ppm

-Incompatible/Reactive

with Magnesium,

Titanium, Aluminum

Methane

-Odourless (<

5000 ppm)

-Colourless

TLV of 1000 ppm -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%


287°C

Reactive with oxidizing

materials


102

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

-Lower flammable limit of 15.4%

-Upper explosive limit of 25%

-Auto ignition temperature

651°C.

-Ignites in presence of oxidizing

materials


materials

-Incompatible with

halogens and acids

Hydrogen

Chloride

-Colorless

-Acidic odour

-Causes severe burns to

exposed skin

-Extremely toxic to

respiratory tract system

-TLV of 2 ppm

-LD50 of 4701 mg/kg.

-Reactive with acetic

anhydride, propylene

oxide, sodium

hydroxide

Chlorine

Gas

-Greenish-

yellow 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

Hydrogen

Sulphide

-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

-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

Ethanol

-Inhalation or contact with

ethanol can cause skin and

eye irritation

-PEL of ethanol is 1000

ppm

-LD50 of ethanol is

3450(oral, mouse).

-Highly flammable and may form

explosive mixtures with air

-Lower flammable limit: of 3.3%

-Upper flammable limit:19%


423°C


material

Acetic acid

-Corrosive

-Causes severe irritation to

the eyes and severe skin

burns

Sulphur -Yellow color -PEL of acetic acid is 10

ppm

-Sulphur dust suspended in air is

readily ignited by flame

-Readily ignites with air


103

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 put 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


104

flow rates, temperature and/or pressure. A more detailed safety analysis of the distillation

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 equipment 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


105

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


106

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 to 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, any 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


107

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.


108

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]


109

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


110

Appendix 2- Detailed Equipment List

Figure A2.1. Detailed equipment list.

Displayed Text Type Specific Extra UtilityB-101 Conveyer Belt

B-102 Conveyer Belt

B-103 Conveyer Belt

B-104 Screw Belt

B-105 Screw Belt

C-101 Chute Gravity Chute Insulation tip None

E-101 Cooling syngas from R-101 Cooling Water

E-102 Cooling Water

E-103 Cooling Water

E-104 Cooling recycle from T-101 before S-101 Cooling Water

E-105 Reboiler for T-101 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 Cooling Water

GC-101 Cyclone None None

GC-102 A Adsorption Column Adsorbent (ZnO) Oxygen

GC-102 B Adsorption Column Adsorbent (ZnO) Oxygen

GC-103 Wet Scrubber Spray Nozzle Cooling Water

P-102 A Pump broth from R-102 to T-101

P-101 A

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

P-105 A

P-105 B

P-106 A

P-106 B

P-107 A

P-107 B

PT-101 Hammer Mill None Electricity

PT-102 Spray Dryer Conveyer Belt Steam

R-101 Gasifier Fluidized Bed (Quartz-Sand) Oxygen/Natural Gas

R-102 Fermenter Bacteria (Clostridium ljungdahlii) Electricity/Water

S-101 Medium Storage None None

S-102 Column feed storage None None

T-101 Distil lation Column Trays None

T-102 Distil lation Column Trays None

V-101 Condenser vessel for T-101 None None

V-102 Condenser vessel for T-102 None None

Cooling of syngas during gas cleaning

Reactor

Storage

Tower

Vessel

ElectricityRubber

Electricity

Pump new medium into R-102

Pump bottoms from T-101

Pump liquid from V-101

Equipment List

Belt

Heat Exchanger

Gas Cleaning

Pump

Pre-Treatment

Pump bottoms from T-102

Pump liquid from V-102

Shell & Tube

None

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 Cause Consequence Action

Low

Low liquid column flow

- 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

Low vapour column flow

- 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

- 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

High

High liquid column flow

- Faulty level control on fermenter, feeding

excess liquid from reactor

- Reflux too high from temperature control

or condenser level control

- Reduce separation quality

- Can lead to flooding of

condenser vessel

- Can damage pumps

- 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)

High vapour column flow

- 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

- Level alarms placed on the column

- Reflux control loop as preliminary control

to prevent rising vapour flow


112

flow too low - Leads to entrainment

- Tray efficiency reduced

- Increase pressure within

column

No

No liquid column flow

- Broken valve/piping

- Reactor failure/shutdown

- Fermenter flash vessel rupture

- Breakdown of pumps

surrounding column

- Column will not operate,

no separation

- 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

No vapour column flow

- 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

- Control loop placed on the reboiler loop

controlling utility flow in ratio to the feed

flow entering the column

Parameter: Temperature


Low Low column temperature

- 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

- Poor column operation

and product purity directly

influenced by column

temperature

- 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.

High High column temperature

- High vapour flow within the column, or

low liquid flow

- Vapour condenser not working properly

- Can lead to unsafe

conditions and poor column

operation causing product

purity to suffer


113

Parameter: Pressure


Low Low column pressure

- 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

- 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

High High column pressure

- 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

- 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.

References

Abubackar, H., & Veiga, M. et. al., (2011). “Biological conversion of carbon monoxide:

Rich syngas or waste gases to bioethanol.” Biofuels, Bioproducts and Biorefining, 5, 93-

114. doi: 10.1002/bbb.256

Advanced Biofuels USA . (2011). Syngas Fermentation, the third pathway to cellulosic

ethanol. Retrieved from http://advancedbiofuelsusa.info/wp-content/uploads/2011/06/3rd-

Pathway-Revision-FINAL-6-17-11.pdf

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

Retrieved From:

http://www.agr.gc.ca/eng/science-and-innovation/agricultural-practices/climate/future-

outlook/climate-change-scenarios/length-of-growing-season-in-

ontario/?id=1363033977515

Alvira, P., Tomás-Pejó, E., et. al., (2010). “Pretreatment technologies for an efficient

bioethanol production process based on enzymatic hydrolysis: A review,” Bioresour.

Technol., 101(13), pp. 4851–4861.

B.C. Ministry of Finance. (n.d.). How the Carbon Tax Works. Retrieved from:

http://www.fin.gov.bc.ca/tbs/tp/climate/A4.htm

Basu, P. (2013). “Biomass Gasification, Pyrolysis, and Torrefaction: Practical Design and

Theory” (2nd ed.). Amsterdam: Elsevier; Academic press.

Canada Revenue Agency. (2015) Classes of Depreciable Property. Retrieved from:

http://www.cra-arc.gc.ca/tx/bsnss/tpcs/slprtnr/rprtng/cptl/dprcbl-eng.html#Class_43

Daniell, J., & Köpke, M. et. al., (2012). “Commercial Biomass Syngas Fermentation.”

Energies, 5, 5372-5417. doi: 10.3390/en5125372

Dunn, J.B., Mueller, S., et. al., (2013) “Land-use change and greenhouse gas emissions

from corn and cellulosic ethanol.”

Environment Canada. (2015). 1981-2010 Climate Normals & Averages. Retrieved from:

http://climate.weather.gc.ca/climate_normals/index_e.html

http://advancedbiofuelsusa.info/wp-content/uploads/2011/06/3rd-Pathway-Revision-FINAL-6-17-11.pdf

http://advancedbiofuelsusa.info/wp-content/uploads/2011/06/3rd-Pathway-Revision-FINAL-6-17-11.pdf

http://www.agr.gc.ca/eng/science-and-innovation/agricultural-practices/climate/future-outlook/climate-change-scenarios/length-of-growing-season-in-ontario/?id=1363033977515



http://www.fin.gov.bc.ca/tbs/tp/climate/A4.htm

http://www.cra-arc.gc.ca/tx/bsnss/tpcs/slprtnr/rprtng/cptl/dprcbl-eng.html#Class_43

http://climate.weather.gc.ca/climate_normals/index_e.html


116

Foust, T.D., Aden, A., Dutta, A., and Phillips, S. (2009). “An economic and

environmental comparison of a biochemical and a thermochemical lignocellulosic ethanol

conversion processes,” Cellulose, 16(4), pp. 547–565.

Intergovernmental Panel on Climate Change. (2013). Climate Change 2013: The Physical

Science Basis. Retrieved from: https://www.ipcc.ch/report/ar5/wg1/

Khanna, M., Dhungana, B., and Clifton-Brown, J. (2008) “Costs of producing miscanthus

and switchgrass for bioenergy in Illinois,” Biomass Bioenergy, 32(6), pp. 482–493.

doi: 10.1016/j.biombioe.2008.08.008

Kratky, L., & Jirout, T. (2010). “Biomass Size Reduction Machines for Enhancing Biogas

Production”. Chemical Engineering & Technology, 391-399. doi:

10.1002/ceat.201000357

Kwiatkowski, J.R., McAloon, A.J., et. al., (2006). “Modeling the process and costs of fuel

ethanol production by the corn dry-grind process,” Ind. Crops Prod., 23(3), pp. 288–296,

Li, H., & Chen, Q. (2012.). “Evaluation of a biomass drying process using waste heat

from process industries: A case study.” Applied Thermal Engineering, 71-80.

doi: 10.1016/j.applthermaleng.2011.10.009

Liska, A. J. H. S. Yang, et. al., (2009) “Improvements in Life Cycle Energy Efficiency

and Greenhouse Gas Emissions of Corn-Ethanol,” J. Ind. Ecol., 13(1), pp. 58–74. doi:

10.1111/j.1530-9290.2008.105.

Maddipati, P., & Atiyeh, H. et al., (2011). Ethanol production from syngas by

Clostridium strain P11 using corn steep liquor as a nutrient replacement to yeast extract.

Bioresource Technology, 102(11), 6494-6501. doi: 10.1016/j.biortech.2011.03.047

Michel, R., and Rapagnà, S., et. al., (2011). “Catalytic steam gasification of Miscanthus X

giganteus in fluidised bed reactor on olivine based catalysts.” Biomass and Bioenergy, 35,

2650-2658. doi: 10.1016/j.biombioe.2011.02.054

National Resources Canada. (2014). ecoENERGY for Biofuels Program: About the

Incentive. Retrieved from:

http://www.nrcan.gc.ca/energy/alternative-fuels/programs/ecoenergy-biofuels/3591

National Resources Canada. (2013). Ethanol – Benefits. Retrieved from:

https://www.nrcan.gc.ca/energy/alternative-fuels/fuel-facts/ethanol/3499

https://www.ipcc.ch/report/ar5/wg1/

http://www.nrcan.gc.ca/energy/alternative-fuels/programs/ecoenergy-biofuels/3591

https://www.nrcan.gc.ca/energy/alternative-fuels/fuel-facts/ethanol/3499


117

Ontario Ministry of Finance. (2015). Gasoline Tax. Retrieved from:

http://www.fin.gov.on.ca/en/tax/gt/

Phillips, S. (2007). Technoeconomic Analysis of a Lignocellulosic Biomass Indirect

Gasification Process To Make Ethanol via Mixed Alcohols Synthesis. Industrial &

Engineering Chemistry Research, 46(26), 8887-8897. doi: 10.1021/ie071224u

Piccolo, C., & Bezzo, F. (2009). A techno-economic comparison between two

technologies for bioethanol production from lignocellulose. Biomass and Bioenergy,

33(3), 478-491.

Pimentel, D., and Patzek, T., (2008). “Ethanol Production Using Corn, Switchgrass and

Wood; Biodiesel Production Using Soybean, in Biofuels, Solar and Wind as Renewable

Energy Systems.” Ed. Springer Netherlands, pp. 373–394.

Roy, P. (2014). “Life Cycle Assessment of Ethanol produced from Lignocellulosic

Biomass: Techno-economic and Environmental Evaluation.” Ph.D. Thesis. The

University of Guelph: Canada

Sánchez, O.J., and Cardona, C.A. (2008). “Trends in biotechnological production of fuel

ethanol from different feedstocks,” Bioresour. Technol.,99(13), pp. 5270–5295.

Saxena, J., & Tanner, R. (2012). Optimization of a corn steep medium for production of

ethanol from synthesis gas fermentation by Clostridium ragsdalei. World Journal of

Microbiology and Biotechnology, 28(4), 1553-1561. doi: 10.1007/s11274-011-0959-0

Seider, W., & Seader, J. (2004). Product and process design principles: Synthesis,

analysis, and evaluation (2nd ed.). New York: Wiley.

Solomon, B. D., Barnes, J.R, and Halvorsen, K.E, . (2007). “Grain and cellulosic ethanol:

History, economics, and energy policy,” Biomass Bioenergy, 31(6), pp. 416–425.

doi: 10.1016/j.biombioe.2007.01.023

Tanner, R., & Miller, L. et. al., (1993). “Clostridium ljungdahlii sp. nov., an Acetogenic

Species in Clostridial rRNA Homology Group I.” International Journal of Systematic

Bacteriology, 43(2), 232-236.doi: 0020-7713/93/020232-05

Towler, G., & Sinnott, R. (2013). Chemical engineering design principles, practice and

economics of plant and process design, second edition (2nd ed.). Kidlington, Oxford,

U.K.: Butterworth-Heinemann.

http://www.fin.gov.on.ca/en/tax/gt/


118

United Nations Conference on Trade and Development. (2009). The Biofuels Market:

Current Situation and Alternative Scenarios. Retrieved from:

http://unctad.org/en/Docs/ditcbcc20091_en.pdf

U.S Department of Energy. (2014). Alternative Fuels Data Center – Fuel Properties

Comparison. Retrieved from:

http://www.afdc.energy.gov/fuels/fuel_comparison_chart.pdf

U.S Department of Energy. (2015). Alternative Fuel Price Report. Retrieved from:

http://www.afdc.energy.gov/fuels/prices.html

U.S. Environmental Protection Agency. (n.d.) Air Pollution Control Technology Fact

Sheet. Retrieved from http://www.epa.gov/ttncatc1/dir1/fsprytwr.pdf

Wang, M., Wu, M., and Huo, H. (2007). “Life-cycle energy and greenhouse gas emission

impacts of different corn ethanol plant types,” Environ. Res. Lett., 2(2), p. 024001.

Woods, D.R. (1983). Cost Estimation for the Process Industries. McMaster University,

Hamilton: Canada.

Woolcock, P., & Brown, R. (2013). “A review of cleaning technologies for biomass-

derived syngas.” Biomass and Bioenergy, 52, 54-84.

Yang, B., and Wyman C.E . (2008). “Pretreatment: the key to unlocking low-cost

cellulosic ethanol,” Biofuels Bioprod. Biorefining, 2(1), pp. 26–40. doi: 10.1002/bbb.49

http://unctad.org/en/Docs/ditcbcc20091_en.pdf

http://www.afdc.energy.gov/fuels/fuel_comparison_chart.pdf

http://www.afdc.energy.gov/fuels/prices.html

http://www.epa.gov/ttncatc1/dir1/fsprytwr.pdf

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