Biodiesel

ChE 422 Design Report

Continuous Production of Biodiesel

Executive Summary The objective of the proposed design project was to produce biodiesel and make it economically competitive with conventional diesel by utilizing low quality feedstock’s and a solid acid catalyst. A Canadian Climate Change Action Plan was produced in 2000 that will be implemented in 2010 that states conventional diesel must have a 2% blend of biodiesel. This plan will create a Canadian demand of 500,000,000 litres per year. SFA will   produce 17,000,000 litres per year of biodiesel by converting triglycerides from waste cooking oil and greenseed canola in a continuous process to meet Saskatchewan’s demand. The plant consists of three main sections; a canola crushing plant, a biodiesel conversion stage and finally a biodiesel/glycerol separation stage. The canola crushing stage uses a screw press to extract 75% of the oil from the canola seed and a membrane

Acknowledgments We would like to thank Dr. Gordon Hill for his assistance and guidance with this  project, and his patients during the days we abused his open door policy. We would also like to thank Dr. Richard Evitts for his assistance in solving any issues we had with our  HYSYS simulation, Dr. Mehdi Nemati for his assistance with heat exch angers and mixers, and Dr. Ding-Yu Peng for his patients and guidance in our design of the hexane stripper. A great deal of thanks goes Dr. Dalai for giving us the opportunity to work on this project and for his additional guidance throughout the year.

Additional acknowledgments;

•

Zenneth Faye, CEO of Milligan Biotech, for letting us tour their Biodiesel plant located on campus.

•

Blair Harker, Chemical Engineering Student, for helping us with design of our 

Table of Contents Executive Summary ...................................................... ....................................................... i Acknowledgments............................................................................................................... ii Table of Contents .................................................... ........................................................ ... iii List of Figures ......................................................................................................... ............ v List of Tables ..................................................................................................................... vi  Nomenclature .......................................................................................................... .......... vii 1.0 Introduction ....................................................................................................... ............ 1 2.0 Qualitative Process Description .................................................................................... 8 2.1.1 Physical Canola Extraction ................................................................................ 9 2.1.2 Solvent Extraction ........................................................ .................................... 12 2.1.3 Meal Preparation .............................................................................................. 16 2.2 Overview of Biodiesel Production ...................................................................... 18 2.2.1 Gum Removal from Canola Oil ....................................................................... 18 2.2.2 Mixing of Reactants ...................................................... ................................... 19 2.2.3 Conversion to Biodiesel via a Fixed Bed Reactor ........................................... 19 2.2.4 Methanol Recovery .......................................................................................... 20 2.2.5 Separation of Reaction Products ....................................................... ............... 21 3.0 Simulation Results ...................................................................................................... 22 3.1 Process Simulator and Fluid Package Used ........................................................ 22 3.2 Canola Extraction................................................................................................ 24 3.3 Meal Preparation ................................................................................................. 28 3.4 Biodiesel Conversion .................................................... ...................................... 29 3.5 Heat Exchangers and Process Heaters ................................................................ 32 3.6 Simulation Summary .......................................................................................... 34 4.0 Process Alternatives ...................................................... .............................................. 35 4.1 Canola Extractor Using Only the Screw Press ................................................... 35

7.0 Economics ........................................................................................................ ........... 51 7.1 Alternative 1: Negating Glycerol Sales .............................................................. 53 7.2 Alternative 2: Using Grade 1 Canola as the Feedstock ...................................... 54 7.3 Alternative 3: Using an Evaporator for Hexane Extraction ................................ 56 7.4 Alternative 4: Selling Glycerol at Current Prices ............................................... 57 7.5 Alternative 5: Selling Glycerol at Current Prices using an Evaporator .............. 58 7.6 Alternative 6: Producing Value Added Products from Glycerol ........................ 59 8.0 Conclusions ........................................................ ....................................................... .. 60 9.0 Recommendations .................................................... ................................................... 62 10.0 References ...................................................... ........................................................ ... 63 Appendix A .................................................... ....................................................... ............ 65 A.1. Sample Calculation for Sizing and Cost of Area 100 ....................................... 66 A.2. Sample Calculation for Sizing and Cost of Area’s 200 and 300 ...................... 78 Appendix B ....................................................................................................................... 82 Appendix C ....................................................................................................................... 98 Appendix D .................................................... ....................................................... .......... 138 Appendix E ..................................................................................................................... 139

List of Figures Figure 1.1: Tailpipe Emissions of Biodiesel Relative to Conventional Diesel Figure 1.2: Transesterification Reaction in the Presence of a Base Catalyst Figure 1.3: Esterification Reaction in the Presence of an Acid Catalyst Figure 1.4: Mechanism for the Simultaneous Transesterification and Esterification of  FFA and TGs Figure 2.1.1i: Rotary Steam Cooker to Evaporate Moisture for the Canola Seed Figure 2.1.1ii: Schematic of o Typical Screw Press used for Physical Extraction of  Canola Seed Figure 2.1.2i: Rotocel Type Extractor for Hexane Leaching of Cake Figure 2.1.2ii:  Nanofiltration Membrane Separator Schematic to Recover Hexane Figure 2.1.2iii: Falling Film Evaporator for Hexane Recovery Figure 2.1.3: Marc Desolventer/Toaster Unit to Produce Meal Figure 2.2: PFD for Area 200 – Conversion to Biodiesel Figure 3.2.1: HYSYS PFD of Canola Extraction Facility Figure 3.4.1: HYSYSTM for the Biodiesel Conversion Process Figure 4.2.1: Canola Extraction Facility Using Evaporator Separator (H-150a) Figure 4.3.1: Basic Schematic of Convention Biodiesel Production Utilizing Homogeneous Catalyst in a Batch Reaction Figure 7.1: Cash Flow Analysis Figure 7.1.1: Cash Flow Analysis with Negating Glycerol Figure 7.2.1: Alberta’s Grade 1 Canola Prices Figure 7.2.2: Cash Flow Analysis Using Grade 1 Canola

List of Tables Table 3.2.1: Important Parameters for Physical Extraction of Canola Oil Table 3.2.2: Important Stream Parameters for Chemical Canola Oil Extraction Table 3.4.1: Important Stream Parameters before Biodiesel Conversion Table 3.4.2: Important Stream Parameters after Biodiesel Conversion Table 3.5.1: Parameters of Heat Exchangers Used in Process as Modeled by HYSYSTM Table 3.5.2: Parameters for Process Heaters in Biodiesel Plant Table 3.5.3: Energy Required for the Process as Calculated by HYSYSTM and by the Short-cut Method. Table 4.1.1: Energy Savings Associated with Screw Press Only Operation Table 4.1.2: Comparison of Physical Extraction versus Combined Extraction Table 4.2.1: Cost Comparison between Evaporator and Membrane Separator  Table 4.2.2: Energy Comparison between Membrane and Evaporator Separator  Table 5.0: Chemical Hazard Information Summary Table 6.11.1: Heat Exchanger Size and Cost for Area 100. Table 6.11.2: Area 100 Unit Sizing and Cost Table 6.14: Sizing and Cost Summary of Fixed Bed Reactor and Flash Drum Table 7.5.1: Summary of Economic Alternatives

Nomenclature  A-Area ABD – Annual depreciation ADME – Direct manufacturing expenses AGE – General expenses AI - Investment AIME –Indirect manufacturing expenses AIT – Income taxes AME – Manufacturing expenses A NCI – Net cash income A NNP – Net annual profit after taxes A NP – Net annual profit AS- Sales ASTM-American Society for Testing and Materials ATE- Total expenses BI – Business interruption loss C bm – Bare module cost CFC – Fixed capital

(m2)

 Le-Length

(m)

m=Mass flowrate

(kg/s)

mbiodiesel-Mss flowrate of biodiesel

(kg/s)

mglycerol-Mass flowrate of glycerol

(kg/s)

nov-Number of stages  NPV - Net Present Value

 N separators –Number of Separators PFD-Process Flow Diagram  ppm-parts per million

P-Power

(kW)

Q-Volumetric Flowrate

(m3/s)

T -Temperature

(°C)

TG - triglycerides

U -Heat Transfer Coefficient

(kW/m2K)

u-Velcoity

(m/s)

V -Volume

(m3)

 X n-Mass/mole ratio in nth stage of underflow/liquid steam  xn-Mole fraction in liquid stream in stripper  Y n-Mass/mole ratio in nth stage of overflow/gas stream

1.0 Introduction Sustainable fuel alternatives have recently become a high priority for many countries and will play a large role in the chemical industry in the near future. Biodiesel is an alternative fuel for the following reasons; it is biodegradable, non-toxic, has low emission profiles and is environmentally beneficial (Krawczyk, 1996). Biodiesel fuel has the potential to reduce the level of pollutants and the level of potential or probable carcinogens (Krawczyk, 1996). Furthermore, biodiesel is an alternative fuel that can be used in unmodified engines with the current fuelling infrastructure (US Dep. Of Energy, 2001) and it is made from renewable biological sources such as vegetable oils and animal fats. With recent technology, diesel engines have become more powerful and 30-35% more efficient than a gasoline engine because of a higher compression ratio. Through the higher compression ratio, diesel engines operate at a higher efficiency, which increases fuel economy by 40%. In addition, diesel fuel contains approximately 15% more energy

 because it burns at a higher temperature. However, fuel additives or a catalytic converter  can be used to reduce nitrous oxide emissions.

Figure 1.1: Tailpipe emissions of Biodiesel relative to Conventional Diesel http://www.ucsusa.org/assets/images/trucks_and_buses/biodiesel_graph.gif 

Biodiesel is the only alternative fuel to possess an overall positive life cycle assessment. NRCan analyzed the life cycle of biodiesel on a based on GHG emissions. Their analysis found biodiesel produces 64-92% fewer emissions compared with

standards which will implement a 2% biodiesel content in diesel by 2012, resulting in a Canadian biodiesel demand of 500 million liters. Biodiesel has not become a vastly popular alternative fuel worldwide due to its higher cost when compared with traditional petroleum diesel. Currently, the National Biodiesel Board states that a 20% blend of biodiesel costs $0.20 more per gallon than  pure petroleum diesel. The two major hurdles in the commercialization of biodiesel are its large production costs and raw material costs. The production of the alternative fuel consists of alkyl esters derived from either  the transesterification of triglycerides (TGs) or the esterification of free fatty acids (FFAs) with short-chained alcohols. The chemical reactions for both transesterification and esterification as shown below.

CH3OCOR 1 CHOCOR 2 CH2OCOR 3

CH3OCOR 1

CH2OH + 3CH3OH

CHOH CH2OH

+

CH3OCOR 2 CH3OCOR 3

the following report. WCO is a preferable feedstock over refined vegetable oil because it is a cheap, low quality oil and will aid in achieving economical feasibility in biodiesel  production. The amount of WCO generated in each country varies depending on the use of vegetable oil. An estimate of the potential amount of WCO from the collection in the European Union (EU) is approximately 0.7 to 1.0 Mtonnes per year . The United States and Canada produce on average, 9 and 8 pounds of yellow grease respectively, per person (Kulkarni, 2006). Currently, WCO is collected from households and restaurants are disposed of as either animal feed or an environmental pollutant. Thus, WCO offers significant potential as an alternative low–cost biodiesel feedstock which could partly decrease the dependency on petroleum-based fuel. The production of biodiesel from WCO is challenging due to the presence of  undesirable components such as free fatty acids (FFAs) and water. Usage of  homogeneous alkali catalyst for transesterification of such feedstock suffers from serious limitation of formation of undesirable side reactions. One such reaction is saponification,

triglyceride (TG) portion of the oil reacts with methanol and base catalyst (usually sodium or potassium hydroxide) to form ester and glycerol. The current process is not economical as it involves a number of steps including washing of the esters to remove acid/alkali catalysts in addition to creating contaminated water disposal issues. Solid acid catalysts have the strong potential to replace liquid acids, eliminating separation, corrosion and environmental problems. Through a summer research project, which this design project was spawned, it was discovered that solid acid catalyst was capable of producing biodiesel to meet ASTM standards in a small scale batch reactor. More specifically, research found Zinc Ethanoate supported on silica is capable of  simultaneous transeterification and esterification to produce quality biodiesel (mechanism can be seen in Figure 4). To our knowledge, there are no reports on the utilization of solid acid catalysts for the production of biodiesel from WCO in a single step. Therefore, in an attempt to develop a robust solid acid catalyst that can simultaneously catalyze esterification as well as transesterification reaction, different types of solid acid catalysts

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C+

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+

C

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CH3

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ESTERS L

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ESTERIFICATION H 3C

+

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H

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R '

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TRANSESTRIFICATION L

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+

OHR"

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+ R 

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L + = acid site on catalyst surface R = alkyl group of acid R ' = c a r b o n c h a i n o f f a t ty a c i d R " = g l y c e r i d e

Figure 1.4: Mechanism for the simultaneous transesterification and esterification of FFA and TGs The use of cheap low quality feed stocks such as waste cooking oil (WCO) in combination with green seed canola oil (canola that has not matured due to frost) will help in improving the economical feasibility of biodiesel. In Saskatchewan, the

water content (no washing of finished product to remove catalyst), and minimal methanol content (methanol recovery maximized by distillation of entire product mixture before separation). There is much debate over the value and demand of glycerol. However, the demand for glycerol is high in the United States, China, and many developing countries. Glycerol has thousands of uses in many industries, such as food, medical, lubrication, and tobacco industries. The use of solid acid catalysts benefits the economics of the process further by reducing the amount of process steps, eliminating water use and toxic waste streams. This  project analyzed the design and economics of continuous biodiesel production utilizing a solid acid catalyst with the above mentioned feedstock and incorporates innovative methods of an onsite canola crusher.

2.0 Qualitative Process Description Description of converting raw canola seed and pre-treated waste cooking oil into  biodiesel, while creating valuable side products such as meal and glycerol. 2.1Canola Extraction Overview Since the primary method to obtain triglycerides for biodiesel comes from Saskatchewan Canola seeds, a process in order to extract this oil needs to be developed. There are many approaches in extracting the oil out of canola seed. Two examples include physical extraction to obtain the oil, and other uses a combination of physical and solvent extraction. Most biodiesel producers use the physical extraction method as it is inexpensive and easy to operate (Faye 2008). However, the physical extraction method has a disadvantage as it leaves a large portion of oil in the seed, therefore increasing operating expenses as more seeds need to be purchased. Another alternative to acquire canola is to purchase the refined canola from a crushing facility already in operation. However, purchasing refined canola oil is

the seed, the remaining material can be sold as a high value meal to livestock owners (O’Brien 2000). The livestock owners demand a feed that is low in fats because any fat   present in the meal will lower the daily intake of solid feed matter by the livestock as explained by Litherland (2005). According to Racz, as long as the fat/oil content in the meal is lower than 6% by weight there should be no decrease in solid material consumption for the livestock (2004). Given this constraint of low oil content animal meal, SFA must meet these requirements through physical and solvent extraction. According to O’Brien, this can be achieved through chemical and physical extraction as there is “<1% residual oil in extracted material” which meets the criteria set above. Even though chemical extraction is expensive, there is an added bonus of a value-added meal. In fact, according to O’Brien “[approximately] 70% of the returns in processing soybeans is due to the sale of meal.” (2000). This is an added benefit to our extraction and biodiesel process as it will give SFA an extra cash flow to help offset the high processing costs of biodiesel

drawback to de-hulling. This oil loss is minimal in comparison to what is gained from the revenue of high protein meal and energy savings. The flaked seed is then sent to a rotary steam cooker to reduce the water content of the seeds down to 9-10% by weight (O’Brien 2000). The water content has to be  brought to this low level because the hexane used in the extraction and the water in the seed are immiscible. Therefore, a seed with a high water concentration prevents the  penetration of hexane into the seed (Perry’s, 1999). This means the effectiveness of the oil recovery will be reduced if the moisture in the seed is not removed. Therefore, a rotary steam cooker, as shown in Figure 2.1.1i, will be employed to cook the seeds, and evaporate the water until the seeds moisture content is at the desired level. According the Perry’s, this type of cooker is “the most common type of indirect heat rotary dryer” (2008) used in industry and thus SFA decided to use this equipment  because of the industry confidence in the said equipment.

  Next, the seeds are sent to a screw press where 75% of the extracted oil is recovered by physical extraction (O’Brien 2000). Figure 2.1.1ii shows the inner workings of a screw press that is utilized in SFA’s canola extraction facility. A screw press is essentially a continuous screw auger that applies pressure on the canola seed as it is conveyed through the screw. The combination of the barrel and the screw form a pocket that is ever decreasing down the length of the screw. The decreasing size of an individual pocket and the friction caused by the moving screw increases the  pressure on the canola seed. This increased pressure causes the oil in the flaked seed to be separated from the canola cake. The separated oil is then collected at the bottom of the screw as shown in figure 2.1.1ii, on the next page. The remaining cake travels down the screw press and collects at the nozzle on the same plane of the screw. The cake is then sent to solvent extraction to recover 95-99% of the in-situ oil in the seed (O’Brien 2000).

2.1.2 Solvent Extraction After the physical extraction, the canola cake is then sent to the solvent extraction train which uses hexane to leach the oil out of the cake. Hexane was chosen as the desired solvent because it has “high oil solubility, ease of evaporation, abundant availability, and historically low price” (O’Brien 1999). Hexane has one major drawback as it can be harmful to worker safety. According to the MSDS for hexane, health effects may include: impaired fertility, harmful through inhalation, irritant, CNS depression, and serious health damage. (MSDS 2005). Solvent extraction begins with canola cake being sent to a rotocel type extractor  as shown in Figure 2.1.2i below. In this extractor, the canola meal is feed in a countercurrent manner to the hexane solvent to maximize the mass transfer of oil into the hexane solvent. Maximum mass transfer is achieved in counter-current operation because the greatest oil concentration gradient in each stage exists when the solvent and the cake are feed in opposite directions from one another.

In the rotocel, solvent is fed into isolated compartments which contain a fixed amount of solid cake. Mass-transfer of solute to solvent is achieved by percolating the solvent through the solid in an open screened bottom compartment. After percolating through the solid, the solute rich solvent runs through a mesh screen into a holding tank    before being pumped into the next stage. Each compartment is rotated around a pivot   point. The solids are subjected at each stage to the solvent spray collected from the  bottoms of the following stage. After the cake has been subjected to the specified number  of stages of solvent washings, the floor of the compartment opens up and the solids are collected and sent for further processing. The cake that was leached is now referred to as marc which consists of 30% hexane and 70% meal (O’Brien 2000). Overflow from the leaching process called micella is loaded with 25% oil and 75% hexane (O’Brien 2000). After the leaching unit, the oil in the micella needs to be separated from the hexane and recombined with the oil recovered in physical extraction. The hexane concentration needs to be reduced down to 200 ppm (O’Brien 2000). The low

Figure 2.1.2ii: Nanofiltration Membrane Separator Schematic to Recover Hexane Taken from Perry’s Handbook (1997), Pg. 22-41, Fig 22-53  Nanofiltration was selected based on the research of Ribeiro et al. They did a study using different types of membrane separators on typical micella compositions (1:3 w/w) in order to recover hexane from the micella (2006).

(2006). This gives comparable results to the evaporator case that concentrates the oil to 60% wt (O’Brien 2000). Given the similar performance between the membrane separator  and the evaporator, the membrane separator was chosen for our design because of the added energy savings. The permeate hexane is recycled to the leaching unit and the retentate micella is sent for second stage separation. The second stage of separation uses steam evaporation to recover more of the hexane from the micella. Using a long vertical tube falling film evaporator as shown in figure 2.2.2iii on the next page, the oil in the micella will be concentrated to 95% wt. The evaporator consists of a vertical single-pass shell and tube heat exchanger  discharging to a very small vapour head. Micella enters into the bottom of the heat exchanger and comes in contact with countercurrent saturated steam. As the steam condenses, the hexane is vaporized from the micella and the oil stays in liquid form. The two phase mixture is sent to a tank where hexane vapour is recovered off the top of the tank and the liquid oil falls to the bottom of the tank where it is drawn off. The

Figure 2.1.2iii: Falling Film Evaporator for Hexane Recovery Obtained from: < http://www.praj.net > The third and final stage of hexane recovery is the counter current hexane stripper  which uses high temperature steam at 130°C in contact with the oil to remove the hexane. After the stripper column, the oil contains less than 200 ppm of hexane and is recombined with the physically extracted oil. This complete process recovers 99% of the oil present in the seeds.

desolventer and toaster (DT) unit. The DT unit consists of hollow stay bolt tray uppers and solid dryer trays to remove the hexane from the meal. Figure 2.1.3 shown below displays a schematic diagram of a typical DT unit that would be used in industry and for  our design.

Figure 2.1.3: Marc Desolventer/Toaster Unit to Produce Meal Taken from :< http://www.fao.org/docrep/t0532e/t0532e25.gif >

2.2 Overview of Biodiesel Production The canola oil obtained from the canola extraction process previously described is combined with citric acid and allowed to settle in gravity separators to remove the gums. From here, it is then combined with 2 million L/year of pre-treated waste cooking oil and 12million L/year of excess methanol. This mixture will be pumped to a pressure of 4240 kPa (g) via reciprocal pump and heated to 200ºC via two heat exchangers and a fired heater. The high pressure, high temperature feed is then sent to a fixed bed conversion reactor that is filled to 50% volume with Zinc Ethanoate/Silica. Within the reactor, simultaneous esterification and transesterification of the TG’s (canola oil) and FFA’s (waste cooking oil) occur to achieve a minimum conversion of 95% to fatty acid methyl esters (biodiesel) and glycerol. The effluent mixture is then depressurized to atmospheric   pressure while maintaining the high temperature. This enables the use of a flash drum with a mist eliminator to capture the methanol as vapor from the product mixture and recycle it leaving a mixture of desirable products: biodiesel and glycerol. The separated

off using a wet vacuum. The gum will then be re-distributed to the canola meal for added  protein. Oil will be drained from the bottom of the tank through a strainer to purify the oil and before being sent to the reactor.

2.2.2 Mixing of Reactants The treated canola oil from SFA’s on-site canola crushing plant is fed to mixer L210 to combine the canola oil, pretreated waste cooking oil and methanol. It should be noted that the waste cooking is treated at both the Regina and Saskatoon processing  plants prior to reception. WCO will be trucked from both processing plants to SFA and   pumped into storage tanks which provide enough oil for two weeks of operation. The waste cooking oil is filtered to remove particle matter and treated to remove the water  created during the frying process. Furthermore, the methanol fed into the second mixer is a combination of fresh methanol and make-up methanol recovered from the flash drum to  be discussed in section 2.2.3.

2.2.3 Conversion to Biodiesel via a Fixed Bed Reactor

methyl esters in the presence of a 3 wt% catalyst and a 1:18 oil to alcohol molar ratio. Furthermore, the catalyst was found to be stable and reusable. Two reactors will be employed by SFA. Catalyst recharging will be performed once a week by switching the continuous flow to the first reactor to the second reactor. The first reactor will be drained and recharged using methanol and hexane. The methanol will remove the polar constituents from the catalyst, where the hexane will remove the non-polar constituents. During this time, an automated system will allow for continuous  production to continue in the second reactor.

  bed reactor (E-213B). The methanol recovery stream is further cooled in another shell and tube heat exchanger using cooling water (E-232). Heat exchanger (213A) removes heat from the biodiesel/glycerol mixture and adds the energy to the reactant stream before the reactor. Finally, the biodiesel/glycerol mixture is sent to gravity settlers for  separation.

2.2.5 Separation of Reaction Products The treated product mixture of biodiesel and glycerol are fed into gravity settlers after being cooled in a heat exchanger in contact with the reaction mixture. Unit H-300 are three gravity settling tanks which will be utilized to ensure continuous production. Once the tank is full it will be left alone for a minimum of 4 hours for sufficient separation to occur. Since glycerol has a higher density of 1215kg/m3 and biodiesel has a lighter density of 880kg/m3, the glycerol will all settle to the bottom and the biodiesel will float to the top. Once separation occurs the tank will be drained off the bottom

3.0 Simulation Results 3.1 Process Simulator and Fluid Package Used HYSYSTM 2006 was the simulator chosen to model the biodiesel process. This   process simulator was chosen because of its capability to incorporate thermodynamic models to predict the energy requirements of the individual units. This allowed us to analyze the energy requirements/surpluses of the process and to optimize the system accordingly. A particular advantage of the energy parameters within the HYSYSTM model allowed us to design heat exchange systems to maximize the utilization of sensible heat within the process. The fluid package used in this simulation was the Non-Random Two Liquid (NRTL) thermodynamic modeling system. The NRTL thermodynamic model was chosen as it was the same model used by Zheng et al in 2003 for modeling an alkali catalyst for   biodiesel production. The NTRL fluid package had adequate modeling parameters for each of the

associated with using HYSYSTM as a process simulator, is that it does not take into account pipeline pressure losses, as there are no pipeline connections for the simulator. Therefore, before an actual facility is completed, a pressure loss survey should be conducted to analyze the pipeline friction losses throughout the plant. From this study it may be determined that the pumping requirements are drastically different than what is  present in the HYSYSTM simulation.

3.2 Canola Extraction The Process Flow Diagram (PFD) of the Canola Extraction process and meal  preparation process is shown in figure 3.2.1 below.

HYSYSTM as it would be a redundant unit in the simulation as HYSYSTM does not model solid crushing. The crushed seed is then sent to the rotary steam cooker, E-111, to bring the water content down to 4% (O’Brien 2000) by heating it to 91°C as shown in Table 3.2.1. In order to model this unit, SFA used a shell and tube heat exchange model in HYSYSTM to determine the amount of low pressure steam that is required to heat up the incoming feed to 91°C. The steam rate was found to be 265 kg/hr at 148°C and 4.5bara.

Table 3.2.1: Important Parameters for Physical Extraction of Canola Oil Stream After roll Mill Steam 1 (Saturated Steam) Steam 2 (Liquid) 2 Oil Recovered Oil to Hexane Recovery

Oil Flowrate (kg/hr) 2080 265 (steam) 265 (water) 2080 1726 354.47

Meal Flowrate (kg/hr) 2080 N/A

2080 0 2080

Temperature °C 25.0 148 147 91 91 91

The counter-current leaching operation using a hexane solvent is modeled as a splitter unit in HYSYSTM . Unfortunately, the splitter was the only unit for leaching simulation as there is no other suitable unit within HYSYSTM. Thus literature data from O’Brien (2000) was used to model and specify the leaching operation of splitter H-120 in the simulation. Staging of the leacher is discussed in appendices A and E of the report. The oil rich micella is pressurized to 15 MPa using a reciprocating positive displacement pump K-122 before being sent to the nanofiltration membrane separators to remove the hexane from the oil. This pressure is needed to accomplish the separation as explained by Ribeiro et al. (2006). The efficiency of the reciprocating pump was specified at 75% as shown in Ulrich (2004). The power requirement of K-122 was determined to be 11 kW from the HYSYSTM simulation. According to HYSYSTM, the discharge temperature of the micella from the pump is 71°C. It will then have to be cooled 59.0°C for membrane separation (Ribeiro, 2006)

only had 8% oil by weight (Ribeiro 2006). Using the research by Ribeiro et al., H-140 splitter was modeled under the same conditions as shown in table 3.2.2 .

Table 3.2.2: Important Stream Parameters for Chemical Canola Oil Extraction Stream Flow Rate Oil wt% Hexane wt% Temperature °C (kg/hr) Micella 1331 36 63 59.0 Micella HP 1331 36 63 59.0 LT* Retentate 436.3 75 25 59.0 Permeate 895.1 82 12 59.0 Hexane Evap 1 103 0 1.0 70.0 Micella Evap 334 98 2 70.0 Hexane Steam 228.0 N/A 98 110.0 Hexane 0.72 N/A 1.0 25.0 make-up 24 1666 9 91 59.0 14 1666 9 91 59.0 *Has a pressure of 15MPa for membrane separation as presented by Ribeiro et al. The oil rich retentate, from the membrane separator, is then sent to an evaporative unit R-150. In this unit, steam is used to evaporate 95% of the hexane originally in the micella feed. As shown in table 3.2.2, the amount of power required to accomplish this separation is 12.5 kW as determined by the Non-Random Two Liquid (NRTL) fluid

in mixer L-161. The hexane rich steam is then cooled to condense the water and hexane using heat exchangers E-164 and E-165. A gravity separator is used to separate both the hexane from the water in unit H-180. Hexane is then recycled back to the hexane leacher  H-120. Hexane recovered from H-140, R-150, D-160, is recycled back to the hexane leaching unit H-120. A make-up hexane stream of 5200 kg/year is needed to maintain the source hexane rate of 1.22x107 kg/year which is needed for solvent extraction. As shown in the hexane recycle stream in figure 3.2.1, a balance was incorporated within the PFD to reduce the degrees of freedom and make the recycle problem solvable. This specification requires that the recycle stream for the leaching system must be equivalent to what was arithmetically determined by mixer L-182. Without this specification, the inlet hexane rate could not be determined in the simulation and the subsequent flows for  the separation would be zero in the simulation.

removed hexane is recycled back into the leaching unit H-120. The meal is cooled to 25°C by contacting the meal to air within the bottom part of the desolventer. Through the use of the air cooled heat exchanger, E-132, the air rate needed for cooling the meal is approximately 424 tonnes/day of air. The resultant meal will be produced at a rate of  2.1tonne/h and sold as a low cost animal feed.

3.4 Biodiesel Conversion After 95%-99% of the canola has been recovered for the feedstock seed, it then enters into the process to produce biodiesel and glycerol. A PFD for the biodiesel conversion process is shown in Figure 3.4.1.

Table 3.4.1: Important Stream Parameters before Biodiesel Conversion Stream Flowrate Wt% Oil Wt% Temperature (kg/hr) Methanol (°C) 32 1887 100 0 40.0 Inlet 1330 0 100 43.0 Methanol 26 (WCO) 148 100 0 25.0 31 3365 60 40 34.0 35 3365 60 40 200.0 Table 3.4.2: Important Stream Parameters after Biodiesel Conversion Stream Flowrate Wt% Wt% Wt% Temperature (kg/hr) Methyl Methanol Glycerol (°C) Ester  Rxn 3365 60 20 20 200.0 Product 33 626 0 100 0 200.0 30 626 0 100 0 63.0 34 2739 75 0 25 200.0 39 2739 75 0 25 55.0 Biodiesel 2064 1.00 0 0 55.0 Glycerol 675 0 0 100 55.0

Pressure (kPa) 101.3 101.3 101.3 4240 4240

Pressure (kPa) 4240 101.3 101.3 101.3 101.3 101.3 101.3

After the canola extraction process, the canola oil is mixed with 208 tonnes/year  of citric acid to emulsify proteins in the canola feed called gums (O’Brien 2000).

According to experimental optimization done by Kathlene Jacobson in the summer of 2008, the reaction occurs best at 200°C. The reactants will be heated to 200°C via two heat exchangers E-213A, E-213B and a fired heater E-213 which was modeled to use 391.7 kW of power. With reactants specified at the correct reaction conditions, they are charged into a  packed bed reactor which is modeled by the conversion reactor P-220. The reaction was specified in the global reaction set as shown in equation X which is the same reaction as specified by Zheng et al (2003).

Triglyceride + 3CH3OH → FAME + Glycerol

(X)

Where the triglyceride is the canola oil and waste cooking oil reactants and FAME is Free Fatty Methyl Esters or biodiesel. Therefore to specify the reaction in the simulation, the stoichiometric coefficients were given as -1,-3,1,1 respectively for reaction X. From Jacobson et al. (2008) the reactor was specified to achieve 95+% conversion using zinc ethanoate/silica. The reactor was modeled as an adiabatic, isothermal and isobaric

residence time will be approximately 3 hours and 100% separation will occur as observed in experiments (Jacobson 2008). The production of biodiesel was calculated by HYSYSTM to be 16,000,000 litres/year, which is within the desired constraints of the design.

3.5 Heat Exchangers and Process Heaters One advantage to our process is over 700 kW can be recovered via 11 heat exchangers as shown in Table 3.5.1.

Table 3.5.1: Parameters of Heat Exchangers Used in Process as Modeled by HYSYSTM Unit Transfer  Energy (kJ/h) ∆Tin ∆Tout ∆Tlm

E-164

E-165

E-124

E-183

E-112

E-121

E-162

E-301

E-213A

E-213B

E-232

3x104

5x105

3x104

6x104

2x105

6x104

2x105

5x105

2x105

1x105

7x105

30 11 19

89 40 61

28 9 17

27 1 8

81 2 21

14 8 11

83 3 24

101 1 22

67 0.75 15

145 0.04 18

60 1 14

Therefore, with the aid of HYSYSTM SFA was able to maximize the recovery of excess heat from various process streams. These HYSYSTM values for the heat exchangers were then compared with the shortcut calculations presented in Ulrich (2004) and were found

Table 3.5.2: Parameters for Process Heaters in Biodiesel Plant Unit No. Description Fuel E-182 Steam Fire-Tube   No. 6 Fuel Oil Boiler  E-213 Thermal Fluid   No. 6 Fuel Oil Heater 

Power (kW) 11 392

As shown in table 3.5.2, HYSYSTM was able to calculate the power requirement of the heaters from the heater simulation unit function. These values were compared to short-cut methods as shown in appendix A and agreed within 1% of each other. It is interesting to note that the process heater used to heat the reactant feed is the most energy intensive unit within the complete process as shown in Table 3.5.3 below. The heater E-213 consumes over 60% of the energy required for the process. Therefore, SFA recommends that more research could be direct towards finding a catalyst that does not require the reactants to  be at a high temperature.

Table 3.5.3: Energy Required for the Process as Calculated by HYSYSTM and by the Short-cut Method. Unit Energy (kW) % of Energy Consumed E-111 156.0 23.9 C-110 2.6 0.4

3.6 Simulation Summary Based on the HYSYSTM simulation, our process is operating with correct mass and energy balances as predicted by the NRTL correlation package. HYSYSTM was an excellent tool used to find the energy requirements for all the units in the process. Based on the simulation of the complete process the utility requirement was determined. The advantage of using HYSYSTM was to configure the heat exchangers to maximize the sensible heat recovery in the process. The simulation also allowed SFA to optimize recycle streams such as; methanol and hexane recycle streams. With only mass and energy balances the recycle and heat recovery would be very difficult to calculate. Therefore, HYSYSTM made our lives easier because it gave us accurate results and made the design easier to optimize.

4.0 Process Alternatives The alternative processes for canola extraction and traditional biodiesel  production will be discussed in this section.

4.1 Canola Extractor Using Only the Screw Press In this alternative, the solvent recovery of canola oil is not considered and the extraction of oil is only achieved by the screw press. The advantage of this alternative is that the energy requirement is minimized. The energy savings associated with this alternative is $55,000 as shown in Table 4.1.1 below.

Table 4.1.1: Energy Savings Associated with Screw Press Only Operation Energy Energy Solvent Cost Screw Cost Energy Extractor  $0.08/kWh  press $0.08/kWh Savings kW $/year kW $/year $/year  254.25 $146,448 158.62 $91,365 $55,083

Table 4.1.2: Comparison of Physical Extraction versus Combined Extraction C bm Solvent

From table 4.1.2, the combined extraction techniques of a screw press and a solvent extraction train will generate a high protein meal worth $5.7 million dollars per  year in revenue. In the first year, the high revenue meal pays for the high capital cost and larger energy requirement in excess of $4.2 million dollars. In other words, the revenue from the meal pays for the hexane extraction facility in one year of operation. The screw   press extraction was considered, but the combined physical and solvent extraction was chosen due to the value added meal.

4.2 Evaporator for Hexane Separation

Another alternative to the crushing process is to use a falling-film evaporator  instead of a membrane separator to perform the initial step in the oil/hexane separation   process. The process is almost identical to the membrane separator cases in terms of  extraction process steps as shown in the HYSYSTM process flow diagram (PFD) in figure 4.2.1 on the previous page. This separator alternative does not affect the biodiesel process in areas 200 and 300, only economics. The main difference between the steam evaporator and the nanofiltration separator is the degree of recovery of hexane from the micella. Steam separation achieves greater relative recovery of hexane at 93% compared to the 87% recovery in the membrane separator (Ribeiro et al. 2006). Also, the evaporator gives a pure hexane recycle stream whereas the membrane separator will give a recycle stream that contains 8.0% wt of oil (Ribeiro et al. 2006). Another advantage to the evaporator is that the individual capital cost for the unit is far less when compared with the membrane separators. The cost of the membrane

The most obvious disadvantage to the evaporator is the energy requirement. The membrane separator will give $13,000 per year of energy savings as shown in Table 4.2.2   below. Given that the plant will run for 10 years, the energy saving per year will be enough to justify the high capital cost of the membrane separators in comparison to the evaporator.

Table 4.2.2: Energy Comparison between Membrane and Evaporator Separator  Energy Energy Cost Evaporator Cost Membrane Cost Savings kW $/year kW $/year  87 $ 14,798 12 $ 2,041 $ 12,757

% Savings 86%

Even though the membrane separator can pay for itself over the lifetime of the   project, it does not give much economic advantage over the steam evaporator. The membrane separator was chosen due to the Kyoto accord and the Canadian Clean Air  Act that requires carbon dioxide (CO2) emissions to be reduced by half of 1990 levels by 2012 (CBC 2006). The carbon credits will have to be monitored closely to determine

needed to convert triglycerides to biodiesel and glycerol in the presence of a short-chain alcohol. However, an acid catalyzed batch reaction is required to neutralize the FFA’s from the waste cooking oil in order to prevent the formation of soap through saponification in an alkali catalytic reaction. This neutralization occurs as an esterification reaction which produces methyl esters and water. The neutralized waste cooking oil can proceed to undergo transesterification by the homogeneous base catalyst. If oil has a free fatty acid (FFA) content greater than 2 wt%, it must be treated with liquid sulphuric acid (homogeneous acid catalyst) to remove the FFAs. WCO   produced from Saskatchewan based processing plants is typically yellow grease that contains 5-15% FFA. Therefore, it must be treated with sulphuric acid to prevent saponification. In the alkali process, sulfuric acid and methanol are mixed together and fed into a   primary batch reactor with pre-treated waste cooking oil to produce methyl esters and water. Excess methanol is added to speed up the rate of reaction of oil conversion to free

second step of transesterification. Once the catalyst and glycerol have been separated from the product mixture, methanol is removed from the biodiesel by washing with water. The water is then removed by drying with a desiccant such as; silica or sodium sulfate. Some processes use distillation to separate the biodiesel, methanol, and glycerol. This method of methanol removal was seen in Milligan Biotech’s facilities on the University of Saskatchewan Campus in October 2007. After separation a quality product of pure biodiesel is obtained. A simplified schematic of the conventional process is displayed in figure 4.3.1 below.

5.0 Safety Analysis In order to identify hazards and operability problems a safety analysis was  performed using a HAZOP study and a DOW Fire and Explosion analysis. The purpose of HAZOP is to determine how the plant might deviate from its designed intent. The Dow Fire and Explosion Index is used to determine the possible explosion hazards and its economical impact on the plant. MSDS sheets were collected and attached in Appendix C. A summary of the MSDS is shown in Table 6 on the next page. The proposed process for the conversion of oil to biodiesel is a fairly safe process that has two major concerns. The first major concern is methanol at high pressure and temperature in the reactor. Since methanol has a low boiling point and a low flash point it has a high material factor of 16 according to the Dow Fire and Explosion Index in appendix C. Since the methanol is heated with a fired heater and pressurized to 4240 kPa creating a very explosive material. The fire and explosion index was calculated to be 243 which is an extremely high value. The full detailed analysis of the fire and explosion

Table 5.0: Chemical Hazard Information Summary LD-50 or LC-50 (route/species)

Component

LEL/UEL

TLV/Exposure Limits

Lower: 1.15% Upper: 7%

176 mg/m TWA – ACGIH 50 ppm TWA-ACGIH 1000 ppm TWA – 500 STEL 3500 ppm TWA – 1760 STEL

Non-flammable

For oil mist: 10 mg/m None known for Liquid

Reaction Hazards

3

Hexane

Canola Oil

C6-H14

Oral toxicity: 25000 mg/kg (rat) Gas toxicity: 48000 mg/kg (rat)

n/a

No known toxicological effects from the product

3

3

Citric Acid

C6H8O7

Oral: 3000 mg/kg (rat)

Oral: 5628 mg/kg (rat) Dermal: 15800 mg/kg

Lower: 0.28 kg/m (dust) 3 Upper: 2.29 kg/m (dust)

No known toxicological effects from this product

Can react vigorously with strong oxidizers

Avoid contact with strong oxidizers Incompatible with oxidizing agents, potassium tartrate, alkali, alkali carbonates and bicarbonates, acetates, sulfides, and metal nitrates

Can react vigorously with oxidizers. Violent reaction with alkyl aluminum salts, acetyl bromide, chloroform +

6.0 Equipment Sizing and Costs This section of the design report gives an overview of the logic behind how each of the units were sized and priced in the process. For a more detailed quantitative description of how each of the units were sized please refer to the sample calculations in appendix A. After the units were sized, and the material of construction was determined, the   bare module cost was found for each units using Ulrich (2004). According to Ulrich (2004), these bare module costs represent the purchase price of the unit, the installation cost, the unit foundation and installing the relative instrumentation. The prices for all the units in the process were obtained from Ulrich using a CE index of 400 as the text was   published in 2004. However, to account for inflation, a more recent CE index of 528.7 obtained for October 2007 from the Chemical Engineering Magazine. (CE index 2008). This index value was applied to all the bare module costs of the units to account for  inflation.

of 1.0 kW (Ulrich 2004). The bare module cost for the roll mill was found to be $42,000 given from Ulrich figure 5.16.

6.2 Rotary Steam Cooker E-111 This unit heats the canola to remove the water in the canola seed to a 1-3% moisture content by heating to 91°C. Using low pressure steam at 148.5°C and the cooker  was modeled as a heat exchanger with a heat transfer coefficient of 700 W/m2K from Table 4-15a in Ulrich. The heat transfer area was found to be 1.6 m2 and the bare module cost for the unit was found to be $15,000.

6.3 Screw Press C-110 One of the most important units of the process is the screw press because the majority of oil is recovered. In order to size this piece of equipment, the power  requirement of the press was determined from Table 4.23c in Ulrich. By pressing 1.2 kg/s of canola seed, the power requirement was calculated to be 2.6 kW. The cost for the screw was determined to be $412,000 using figure 5.58 in Ulrich (2004).

6.5 Membrane Pump K-122 The micella is pressured up to the desired pressure of 15 MPa used for the membrane separators as explained by Ribeiro et al. (2006) using a multistage reciprocating pump. Since the temperature of the pump is operating below 200°C and the   pressure is relatively high, cast steel was used to construct the multistage reciprocating  pump. From the HYSYSTM simulation and from the shaft work equation in Ulrich (2004) the pump requires 11 kW of power in order to achieve the desired outlet pressure. The  bare module cost of the pump was determined to be $162,000 using figures 5.49 to 5.51 in Ulrich (2004).

6.6 Membrane Separators H-140 Using the research put forth by Ribeiro et al. (2006), nanofiltration membranes were designed for the initial hexane separation. By using the maximum experimental flux of 3.0x10-5m3/s.m2 (Ribeiro 2006) the total area of membranes was found to be 22.6m2. Since the maximum surface area of an individual membrane separators is 2.4m2 (Ulrich 2004), 10 parallel separators will have to be employed to handle the volumetric flowrate

6.8 Hexane Stripper D-160 The rigorous calculation method was used to determine the dimensions of the stripper and was completed by using the stage wise equilibrium integration technique as shown in appendix E. The hexane steam equilibrium was modeled using Raoult’s law of  ideal liquid-vapor equilibrium. This approximation gave reasonable results for a   preliminary design of the column (O’Brien 2000). Using the stage wise equilibrium method, the height of the column was determined to be 2.8m. The diameter of the column was determined by finding the stage efficiency in the column achieved by using 6mm Raschig Rings and using that value in equation 4.85 in Ulrich (2004). Using this method, the diameter of the column was found to be 0.44m or 44 cm. The stripping unit will be made from stainless steel because the steam in the column is corrosive. Given this material of construction, the stripper will cost $48,000 as given by figures 5.44-5.47 in Ulrich (2004).

6.9 Hexane Flash Drum H-180 In the flash drum, water and hexane are separated from each other from the

an evaporator was found to have a height of 1.5m and a diameter of 0.7m as calculated from pg 133-136 in Ulrich (2004). The lower portion of the column was found by adding two evaporator heights to the bottom of the evaporator. Therefore, the total height for the desolventer and heater was 2.8m. The cost of the unit was determined to be $100,000 as shown in figure 5.33 for a Rotary Ro tary Dryer in Ulrich (2004).

6.11 Heat Exchangers Since all the heat exchangers are operating at temperature below 200°C, the material used to make the heat exchangers will be carbon steel. Since there are so many heat exchanger units, Table 6.11.1 shows the area and costs associated with each heat exchanger in the process. A sample calculation for these values is given in appendix A for E-124 following the procedure given in Ulrich on pg. 191 (2004) and from Dr.  Nemati’s notes (CHE 324 2006).

Table 6.11.1: Heat Exchanger Size and Cost for Area 100. Bare Module Cost C bm HX Unit Heat Transfer Area m2 (2008) E-112 7.0 $ 11,895.75

Table 6.11.2: Area 100 Unit Sizing and Cost

Unit Nam Name Unit Num Number Roll Mill Cooker

E- 111

Screw Press

C- 110

Leacher

H- 120

Pump Membrane Evaporator Stripper

KK- 122 H- 140 R- 150 D- 160

Height

Diameter

Area

m

m

m

2

Power

Material

kW

$

1

Carbon Steel Carbon Steel

$42,000 $15,000

2.6

Carbon Steel

$412,000

Carbon Steel

$513,000

Cast Steel

$162,000 $132,000 $4,500 $48,000

1.6

10 22.6 0.6 2.8

0.3 0.44

Cos Cost

Carbon Steel Stainless Steel

6.12 Mixer Design First a volumetric flowrate was needed to design the mixer for the separation of  gum from canola oil. The volumetric flow rate going into the mixers was 41m3/day. Three mixing tanks were used and rotated continuously. A plant height restriction of 5m was implemented therefore the unit height was determined to be 4m. 4m. Using a 2:1 height

requirement of 111MJ. Using figure 5.42 in Ulrich the current estimated price was found to be $55,800 for the three units.

6.13 Fixed Bed Reactor The fixed bed reactor is used for the conversion of triglycerides and free fatty acids into biodiesel (methyl esters) and glycerol. Due to the high temperature and  pressure of this vessel, a stainless steel material will be used. The fixed bed reactor was difficult to size because of no kinetic data. Without kinetic data, the size of the reactor  was based on the weight hourly space velocity, as defined by Vartuli et al, to obtain the residence time of the liquid. Furthermore the reactor was sized based on 50% volume of  catalyst. From laboratory work, it was determined that a 3 wt% catalyst was capable of  achieving the desired conversion. The volume of the reactor was found to be 175 L with a liquid residence time of 1.3min. The bare modular cost of this particular fixed bed reactor  is $57,600. The remainder of the reactor parameters can be viewed in table 6.14 on the following page.

Table 6.14: Sizing and Cost Summary of Fixed Bed Reactor and Flash Drum Unit Number

P 220

D 230

Fixed Bed Reactor

Flash Drum

Height of Unit (m)

0.96

2.04

Diameter of Unit (m)

0.48

1.94

Residence time (min)

1.32

10

Volume (m^3)

0.173

12.06

$57,600.00

$144,900.00

Description

Cbm

‐

‐

6.15 Gravity Separators To design the gravity separators for this plant the first variable needed was the inlet flow rate of glycerol and biodiesel. The combined flowrate was 71.5m3/day and 3 tanks were designed to hold this volume for 4 to 6 hours of separation time. The dimensions of the tanks were 5m for the height and 2.5m for the diameter, giving a total volume of 24.5m3. Table 5.61 in Ulrich was used to determine the bare module cost of  $1,100 for each unit.

7.0 Economics A complete economic analysis proves that SFA has a strong potential for   profitability. Many alternatives were considered, and the plant always sustained enough income to generate a net positive cash flow. A detailed cost analysis was performed for the alternative using a membrane separator for hexane extraction and conservative estimates for the price of diesel, glycerol and meal at $80/barrel, $.4/kg and $.38/kg respectively. The revenue generated from this  plant was estimated to be 16.4 million dollars. The biodiesel, the meal and the glycerol  produced 8.8, 5.7 and 1.9 million dollars respectively. The total capital investment for the  proposed biodiesel plant was estimated to be $3,300,000. The total manufacturing costs was estimated to be $14,700,000 per year as shown in the detailed summary in table B-1 in appendix B. The direct manufacturing expense was estimated to be 13.8 million dollars per  year. The cost for the raw materials, mainly greenseed canola, was estimated to be 11.7

general expense was estimated to be $640,000. Depreciation was determined to  be$290,000 annually and income tax was estimated to be $600,000 per year. From all of  these estimates, the after tax rate of return was determined to be a promising 48.6% as shown in figure 7.1 below. The discounted breakeven point (DBEP) was 3.5 years; the undiscounted and 10% discounted net present values were estimated to be $11.1 and $5.3 million dollars after 10 years of operation. In comparing this scenario with the one presented in alternative 7.3 it can be seen that the evaporator is more economical after ten years of  operation. However, the membrane separator was chosen because it produces less CO2 emissions and requires less energy.

7.1 Alternative 1: Negating Glycerol Sales The first alternative assumes the glycerol produced in plant is worth nothing. In order  to produce a net positive cash flow, the price of biodiesel has to be increased to $90 per   barrel, which is still realistic because the current price of diesel is $100 per barrel. This method produces a discounted cash flow rate of return of 30%. The undiscounted and 10% discounted net present values were determined to be $6,000,000 and $2,500,000 respectively, after 10 years of operation. The discounted break even point occurs after 4.5 years of operation. This alternative is the most probable worst case scenario and the plant still generates a positive net present worth.

7.2 Alternative 2: Using Grade 1 Canola as the Feedstock

The second alternative is the worst case scenario because the plant would have to  purchase expensive grade 1 canola seed as the major feedstock instead of the inexpensive greenseed canola. Since canola seed is the main raw material, the fluctuations in the price will have an extreme effect on our economics. The price of canola used for the analysis was $650/tonne.

and $950,000 respectively, after 10 years of operation. The discounted breakeven point will occur at 6.5 years.

Figure 7.2.2: Cash Flow Analysis Using Grade 1 Canola

7.3 Alternative 3: Using an Evaporator for Hexane Extraction

The third alternative studied would be to use an evaporator instead of a membrane separator for the hexane extraction unit. This method produces a discounted cash flow rate of return of 50%. The undiscounted and 10% discounted net present values are estimated at $11,500,000 and $5,500,000 respectively, after 10 years of operation. The discounted breakeven point requires 3.1 years of operation. Although this alternative  produces a higher discounted cash flow rate of return in comparison to the membrane case. Even though the evaporator case has favorable economics, it produces more CO2 emissions and consumes more energy than the membrane separator. Therefore, the membrane separator will be used for our project design.

7.4 Alternative 4: Selling Glycerol at Current Prices The fourth alternative was estimated using the current price of glycerol at $.7/kg. This alternative produces a discounted cash flow rate of return of 87.5%. The undiscounted and 10% discounted net present values are estimated at $22,800,000 and $11,800,000 respectively, after 10 years of operation. The discounted breakeven point requires 2.6 years of operation. This is best case alternative, so if glycerol prices remain high this biodiesel plant could make a fortune.

7.5 Alternative 5: Selling Glycerol at Current Prices using an Evaporator

The fifth alternative uses the same prices for the produc ts as in alternative 3, but uses an evaporator instead of the membrane separator for hexane extraction. This alternative  produces a discounted cash flow rate of return of 88.9%. The undiscounted and 10% discounted net present values are estimated at $23,200,000 and $12,100,000 respectively, after 10 years of operation. The discounted breakeven point requires 2.6 years of  operation.

7.6 Alternative 6: Producing Value Added Products from Glycerol Another alternative considered was to produce value added products from the glycerol produced. The glycerol would be sent to a fixed bed reactor with a silica alumina catalyst. The four main products produced from this reaction are acetol, formaldehyde, acrolein and acetaldehyde. The reaction also produces many side products such as; Hydrogen gas, carbon monoxide, carbon dioxide, methane, acetone and allyl alcohol. Since these products have different volatilities the separation process to acquire the four  main products would be too costly. A summary of the alternatives are displayed in a table below.

Table 7.5.1: Summary of Economic Alternatives Alternative

DBEP i=10% (years)

NPV i=0% ($)

NPV i=10% ($)

Rate of Return (%)

Glycerol Worth Nothing

4.5

6,000,000

2,500,000

30

Grade 1 Canola Oil

6.5

3,250,000

950,000

18.5

2.6

23,200,000

12,100,000

88.9

Regular Glycerol price evaporator

8.0 Conclusions 1. SFA will produce 15,000,000kg of biodiesel per year at $.59/kg giving an annual revenue of $9,000,000. 2. SFA will produce 15,200,000kg of canola meal per year at $.38/kg giving an annual revenue of $5,700,000. 3. SFA will produce 4,900,000kg of glycerol per year at $.59/kg giving an annual rate of $1,900,000. 4. Glycerol was deemed to be a viable source of income because the world demand is 1,600,000 tonnes/year and is increasing with new technology. 5. SFA will consume 30,000,000kg of greenseed canola per year at $.33/kg giving an annual cost of $9,800,000. 6. SFA will consume 1,100,000kg of waste cooking oil per year at $.15/kg giving an annual cost of $100,000. 7. SFA will consume 5,100,000kg of methanol per year at $.34/kg giving an annual cost of $1,700,000.

13. The fixed bed reactor was deemed to be a severe hazard and will be isolated from the plant. 14. The heterogeneous acid catalysis process has fewer process steps in comparison with a homogeneous alkali catalyst process. 15. By utilizing a solid catalyst and a low quality feedstock, the biodiesel price is theorized to be competitive with current diesel prices.

9.0 Recommendations 1. The biodiesel reactor needs more kinetic data to obtain more accurate reactor   properties and sizing. 2. Membrane separators need to be tested in pilot facility to determine their  compatibility with the other units in the process. 3. Pipeline pressure survey should be completed before an actual plant is built to more accurately predict frictional losses and pumping requirements. 4. Supercritical CO2 oil extraction should be investigated as an attempt to lower the solvent extraction energy requirements in the given process. 5. More research should be conducted to find a catalyst that requires a lower  reaction temperature. 6. Meal standards may change and physical extraction can be utilized for canola recovery to minimize energy and unit costs associated with solvent extraction. 7. Propylene glycol can be created as a value added product if glycerol demand decreases or glycerol supply increases from other biodiesel projects. 8. Profitability of project is largely dependent of on feedstock prices. More testing should be done on feedstocks other than canola seed as canola prices may rise to a level that will make this process uneconomical. 9. Further study on the characterization of the finished produ ct is needed to ensure

10.0 References Carbonfund, Leading Offset Provider Value Chart, Carbonfund.org. 2007. April 3, 2008. < http://www.carbonfund.org/> CBC News, In Depth Kyoto and Beyond Trading Carbon, CBC News. November 3, 2006. < http://www.cbc.ca/news/background/kyoto/carbon-trading.html> CE Index October 200, Economic Indicators, Chemical Engineering. February 2008. Pg. 68. Faye Zenneth, CEO Milligan Biotechnology, Foam Lake. Personal Communication. Forge, Frederic. Biodiesel – An Energy, Environmental or Agricultural Policy? Library of Parliament. 8 February 2007 Litherland N.B. et al., Dry Matter Intake is Decreased More by Abomasal Infusion of  Unsaturated Free Fatty Acids than by Unsaturated Triglycerides, J. Dairy Sci. 99 (2005) 632-643. Hill Gordon. Chemical Engineering Process and Design I, CHE 325, University of  Saskatchewan (1999). Kiss, A.K.; Dimian, A.C.; Rothenberg, G. Solid Acid Catalysts for Biodiesel Production  – Towards Sustainable Energy.  Adv. Synth. Catal. 2006 , 348, 75-81. Krawczyk, Tom. Biodiesel. INFORM. Vol. 7, No. 8 August 1996 800

Racz V., Christensen D.A. , Whole Canola Seed use and Value, Prairie Feed Resource Centre. (2004). Ribeiro et al., Solvent recovery from soybean oil/hexane miscella by polymeric membranes, J. of Membrane Science 282 (2006) pg. 328-336. Sea News Circular, Canada Moves on Renewable Fuels, Sea News Circular. Vol: X, Issue:10. Jan., 2008. Titipong, I.; Kulkarni, M.G.; Dalai, A.K.; Bakhshi, N.N. Production of Biodiesel from Waste Fryer Grease Using Mixed Methanol/Ethanol System US Department of Energy, Clean Cities – Alternative Fuel Information Series – Fact  Sheet, May 2001 Ulrich G., Vasudevan P. Chemical Engineering Process and Economic A Practical Guide 2nd edition. Process Publishing. Durham. (2004). Zaher F.A., El Kinawy O.S., El Haron Ha ron D.E., Solvent extraction of jojoba oil from pre pressed jojoba meal, Grasas y Aceites. 55, 2, (2004), 129-134. Zheng S., Kates M., Dube M.A, McLean D.D. Biodiesel Production from waste Cooking Oil 1. Process design and technological tech nological assessment, Bioresource Technology 89 (2003) 1-16.

Appendix A

Sample Calculations

A.1. Sample Calculation for Sizing and Cost of Area 100  Note: All Figures Referred to for the Sample Calculations are from Ulrich (2004) 1. Sample Calculation for Power Requirement of Roll Mill .

P = 0.20 m R Equation from Table 4.5a for Roll Crusher of medium hardness P = 0.20(1.16kg / s)(4) P = 0.925kW  Capacity = 1.16 kg / s Cp = $15,000 Fbm = 2.10 Cbm = (Cp)( Fbm)

(CE index2008 ) (CE index 2004 )

Cbm = $15,000 x 2.10 x

528.7 400

Cbm = $41,635.13

2. Sample Calculation for Rotary Steam Cooker E-111 Δ H  =  H 2 −  H 1 For  Steam Change Δ H  = 623 . 16

kJ  kJ  − 2742 . 9 K From Steam Tables kg kg

Δ H  = 2119 . 7

kJ  kg

Supplied  to Coo ker 

.

P = Δ H  m kJ 

2

kg

Heat Transfer Coefficient Obtain from Table 4-15a, for Conde nsing Water on Hot side and Low-Viscosity Organics in Cold Side.

The Bare module cost for the rotary steam cooker was found using Ulrich from Figure 5.24 for evaporators. Since the unit is at temperatures below 200°C, then carbon steel was selected for materials of Construction. Since the steam cooker was operated at atmospheric pressures Fp=1.0

Cp = $5000 Fig 5.24 Fp = 1 Fbm = 2.3 Cbm = Fbm × Fp × Cp ×

(CE index 2008 ) (CE index 2004 )

Cbm = $15,200.13

3. Sample Calculation for the Shell-&-tube Heat Exchanger. E-124 T  − T 2  R = 1  Eq. 4 − 70  from Ulrich t 2 − t 1  R =

(43.00 ° C − 50.04 ° C ) (59.00 ° C − 71.01° C )

 R = 0.59 S=

(t 2 − t 1 ) (T 

)

 Eq. 4 − 70  from Ulrich

The bare module cost associated with heat exchangers was determined by using figures 5.36 to 5.38 in Ulrich (2004). All the heat exchangers will be constructed using carbon steel because all the units will be operating at temperatures below 200°C. A pressure factor of one was used for all heat exchangers because the heat exchangers are all at atmospheric pressure. Cp = $2300.00 Fm = 1 carbon steel Fp = 1 Fbm = 3 Cbm = Cp × Fbm ×

(CE index 2008 ) (CE index 2004 )

Cbm = $9,120.80

4. Sample Calculation for Screw Press Unit C-110 . 0.5

P = 2.5 m

P = 2.5(1.10

kg 0.5 ) s

P = 2.62kW  Duty Equation Obtained from Table 4.23c in Ulrich (2004). Using Figure 5.58, the Bare Module Cost of the screw press was determined. Carbon steel is used to be more economical

5. Sample Calculation for Number of Stages and Residence Time of Solid-Liquid   Extractor as well as the Bare Module Cost H-120 F solid  F solvent  F solid  F solvent 

 X n =

=

0.664 kg / s 0.370 kg / s

= 1.80

F solvent  F solid 

(Y n+1 − Y 1 ) +  X 0 Operating Line

 X 0 = 0.556(0.282 − 0.282) + 0.165  X 0 = 0.165kgSolute / kgSolids Y n*+1 = mX n Y 1* = (3.255)(0.165) Y 1* = 0.537 kgSolute / kgSolvent  (Y 1* − Y 1 ) = 0.537 − 0.282 (Y 1* − Y 1 ) = 0.256 kgSolute / kgSolvent  1 (Y  − Y 1 ) * 1

= 3.91

Y n

nov =

∫  (Y 

Y n +1

* 1

1

− Y 1 )

dY 

nov = 11 stages ( Simpson' s Rule)

6.

Sample Calculation for Power Requirement of Membrane Pump Unit K-122 .

q ΔP P= e (6.79 x10 −4 m 3 / s )(15 x10 7 Pa) P= 0.75 P = 13.6kW  Efficiency was estimated from Table 4.20 in Ulrich for Rec iprocating Pump. Bare Module Cost was determined from figures 5.49 to 5.51. Cast Steel was used as the temperature of the Process was well below 200°C. Cp = 25,000.00

Fm = 1.8 Cast Steel Fp = 1 < 10 b arg Fbm = 4.9 Cbm = Cp × Fbm ×

(CE index 2008 ) (CE index 2004 )

Cbm = $161,914.38 7. Sample Calculation for Size and Number of Nanofiltration Membrane 140 QTotal = 6.79 x10 −4 m 3 / s  A

=

QTotal

=

6.79 x10 −4 m 3 / s

Units H-

Bare Module Cost was obtained by using figure 5.57a in Ulrich. Nanofiltration was used for costing as that was the same membrane used by Ribeiro et al. 2006.

Cbm=$132,175.00

8. Sample Calculation for Rising Film Evaporator Unit R-150 .

m vap = 4 .44 x10 − 2 kg / s

Δ H vap = 334 .5 kJ  / kg .

P = Δ H vap m vap P = (334 . 5 kJ  / kg )( 4 . 44 x10 − 2 kg / s ) P = 12 .47 kW   A jacket  =

U (T s − T ) lm

12 .47 kW  kW  0 . 8 2 (87 . 5 K ) m K  = 0 .253 m 2

 A jacket  =  A jacket 

P

⎛  ρ l −  ρ g  ⎞ ⎟ ⎜  ρ  g ⎟ ⎝   ⎠ ⎛ (781 .3kg / m 3 − 3.06 kg / m 3 ) ⎞ ⎟⎟ u = 0.06⎜⎜ 3 3 . 06 / kg m ⎝   ⎠ u = 0.957 mls u = 0.06⎜

.

 Across − sec tion =

m  ρ g u

⎛   ⎞ 0.117 kg / s ⎟⎟  Across − sec tion = ⎜⎜ 3 ( 3 . 06 / )( 0 . 957 / ) kg m m s ⎝   ⎠  Across − sec tion = 0.040 m 2

⎛ 4 A  ⎞  D = ⎜ cross −sec tion ⎟ π  ⎝   ⎠

1/ 2

⎛ 4(0.040 m 2 ) ⎞ ⎟⎟  D = ⎜⎜ π  ⎝   ⎠  D = 0.23m  H  =  H  =

 A jacket  π  D

1/ 2

+ D

0.253 m 2 π (0.23m )

+ 0.23m

9. Sample Calculation for Hexane Stripper Sizing Using Raschig Rings. Unit D-160 Determine the diameter of the column from short cut methods in Ulrich on Page 240. Figure 4.31.

 x − axis =

mM l

l

 ρ l

⎛ 0.756*100kg / kgmol* 0.0102Pa.s ⎞ ⎟⎟ 3 1150 / kg m ⎝   ⎠

 x − axis = ⎜⎜

 x − axis = 0.00068 es = 0.178 Stage Efficiency for  Bubble− cap Tray  HETP =  HETP =

 L K la 9.21 x10 − 5 kgmole / s 5.844 x10 − 4 kgmole m 2 / m 3 s

 HETP = 0.987 m  D =  D =

 HETP × es

 Equation 4.85  from Ulrich 0 .4 (0.987 m )(0.178) 0 .4

 D = 0.44 m  z =  Nov × HETP

Bare Module Cost for the Stripper was determined using Figures 5.44-5.47. Stainless Steel a there is corrosive materials in the stripper. 6mm Ra schig rings were used in the tower.

C ss p = $3,800.0 Fbm = 1.2 Cbm = $4,560.0 Fp = 1 Fm = 4 a = 9.5 F bm a × Cbm = Cbm × F bm

(CE index 2008 ) (CE index 2004 )

Cbm = $47,715.18 10. Sample Calculation for Desolventor Energy Requirements/Sizing Units H-130 and E132 Treat as an evaporator that removes hexane from Meal. .

m hexane = 2.81 x10 −1 kg / s

Δ H vap = 334.5kJ / kg .

P = Δ H vap m hexane P = (334.5kJ  / kg )( 2.81 x10 −1 kg / s)

⎛  ρ l −  ρ g  ⎞ ⎟ ⎜  ρ g ⎟ ⎝   ⎠ ⎛ 625kg / m 3 − 3.06kg / m 3  ⎞ ⎟⎟ u = 0.06⎜⎜ 3 kg m 3 . 06 / ⎝   ⎠ u = 0.855m / s u = 0.06⎜

.

 Across −sec tion =  Across −sec tion  Across −sec tion

m  ρ g u

⎛   ⎞ 9.49 x10 −1 kg / s ⎟⎟ = ⎜⎜ 3 ( 3 . 06 / )( 0 . 855 / ) kg m m s ⎝   ⎠ 2 = 0.36m

⎛ 4 A  ⎞  D = ⎜ cross−sec tion ⎟ π  ⎝   ⎠

1/ 2

⎛ (4 × (0.36m 2 ) ) ⎞ ⎟⎟  D = ⎜⎜ π  ⎝   ⎠  D = 0.68m  H  =

 A jacket  π  D

1/ 2

+ 3 D

⎛  1.6m 2  ⎞ ⎟⎟ + 3(0.68m)  H  = ⎜⎜ ( 0 . 68 ) π  m ⎝   ⎠  H  = 2.83m

11. Sample Calculation for Hexane Flash Drum H-180

⎛  ρ l − ρ g  ⎞ ⎟ ⎜  ρ g ⎟ ⎝   ⎠ ⎛ 973.0kg / m 3 − 3.06kg / m 3  ⎞ ⎟⎟ = 0.06⎜⎜ 3 kg m 3 . 06 / ⎝   ⎠ = 1.138m / s

u s , g = 0.06⎜ u s,g u s,g

⎛  4VMg  ⎞ ⎟  D = ⎜ ⎜ πρ g u s , g ⎟ ⎝   ⎠

1/ 2

⎛ 4 × 1.6169 x10 −5 kgmol / s × 86.18 kg / kgmol ⎞ ⎟⎟  D = ⎜⎜ 3 kg m m s ( 3 . 066 / )( 1 . 138 / ) π  ⎝   ⎠  D = 0.02m

1/ 2

Design to hold 10 minutes of liquid in the bottom of the Separator  .

V liq =

m  ρ 

(0.167 hours)

m3 V liq = 0.2235 (0.167hour ) h V liq = 3.72 x10 −2 m 3

⎛ 4V   ⎞

Procedure Outlined from Dr. Hill’s CHE 325 notes. Properties for water and hexane obtained from Perry’s Handbook. Bare module cost was determined from figures 5.44 to 5.46 in Ulrich. Carbon Steel was used because there is no corrosive material in the feed.

Cp = $3,000.00 Fm = 1 Fp = 1 a = 4.5 F bm a × Cbm = Cp × F bm

(CE index2008 ) (CE index 2004 )

Cbm = $17,843.63

A.2. Sample Calculation for Sizing and Cost of Area’s 200 and 300 12. Sample Calculation for the sizing of the canola mixers (L-163) Approximate volume needed for each tank 

Using a rule of thumb of a 2:1 height to diameter ratio (Ulrich 579). A height of 4m and a diameter of 2m was used to calculate the actual volume. The tank is cylindrical with a conical funnel at the bottom. The propeller size was determined to be .5m in diameter  which is the average maximum mixing blade size. (Ulrich 578)

From figure 5.42 in Ulrich using 6kW as the power consumption, the purchased equipment cost was estimated at 7,000$ and the unit was made of carbon steel giving a FBM of 2.0 Calculation for the bare module cost

Calculation for the current price using 528.7 as the current CE index value

13. Sample Calculation for Fixed Bed Reactor 

Using the LHSV definition as layed out by Vartuli et. Al as the weight of liquid reactant (ie. Methanol and oil) contacting a given weight of catalyst in one hour, we can find the LHSV based on the flowrate and the weight of catalyst:

 LHSV  =

amount  _ of  _ liquid  2222 L / hr + 1687 L / hr  = = 45.3hr −1 86.3 L Volume _ catalyst 

θ  = residence _ time = 0.022hour  = 1.3 min utes

We then know the reactor volume can be calculated using the volumetric flowrate of the oil multiplied by the residence time plus the volume of the catalyst.

V  = (2222 L / hr + 1687 L / hr )0.022hr −1 + 86.3 L = 1

The diameter and height of the reactor can then be calculated using the rule of thumb that the height should be twice the diameter. 2

3

π d  ⎛ d  ⎞ V  = π r  d  = π ⎜ ⎟ (2d ) = 2 ⎝  2 ⎠ 2

1/ 3

⎛ 2V  ⎞ d  = ⎜ ⎟ = 0.48m ⎝  π   ⎠  H  = 2d  = 0.96m 14. Sample Calculation for Fixed Bed Reactor 

15. Sample Calculation for the sizing of the three gravity separators (H-300). Calculation for the approximate amount of glycerol/biodiesel mixture for each tank.

Calculation for the actual volume of each tank. The heights of the tanks were 5m and the diameters were 2.5m. The tank was cylindrical with a conical drain at the bottom.

Using figure 5.61 in Ulrich and a volume of 24.9m3 the capital cost was estimated at

Appendix B Economics Please Refer to Disc for More Details

83

84

86

87

89

91

93

95

97

Appendix C

HAZOP, MSDS, DOW FIRE and Explosion HAZOP/Safety Considerations 1: Identify Hazards by considering the following Process Parameters : i)

Flow: - design for a no flow situation (pump cavitation/failure) and purchase (2) secondary pumps for the biodiesel system - Flowrate to reactor is relatively low, and should not be considered hazardous

ii) Time: -Because flowrates are low and residence time in the reactor is short, overflow

vii) Viscosity: liquid is used, no slurries or solids, no chance of freezing. Product liquid of  high viscosity, thus pipe fowling should be considered to prevent build-up of   pressure viii) Temperature: - Reactor operates at high temperature of 200 °C and thus stainless steel will be utilized ix) pH: -  N/A x)

Separation: - Occurs in the gravity settling tanks after the reactor 

xi) Level: - Flowrates and residence time are relatively low, therefore overflow is not of  concern xii) Speed: - Pumps have moving parts, consider operator maintenance manuals for the  pump attached to the reactor  xiii) Information: - Documentation such as operating manuals, equipment specifications, design criteria and MSDS should be supplied to new users to communicate details of   plant

General Operating Problems: xvi) Desired conversion is not achieved.: - If the catalysts activity degrades, then the necessary conversion of 95% will not be obtained. - Consider purging the reactor with methanol and hexane to recharge the catalyst (methanol to remove polar particles from the catalyst and hex ane to remove non-polar particles) or to replace the catalyst since it is silica based and low-cost xvii) Pressure build-up: - Because the reactor is operating at a high pressure in order to keep the methanol in it’s liquid phase at 200°C, pressure drop and build-up is of main concern - Pressure drop should be carefully monitored as well as the pumps  performance xviii) Piping Design2: High viscosity may cause fowling of pipes - “advisable to: 1) Avoid the use of screwed fittings whenever practical. 2) Use welded fittings with long radius ells; avoid tees when possible.

Plant Safety1: i) Safe Design: See HAZOP above ii) Pollution Prevention: - When catalyst is replaced, proper disposal methods will be required. - If spills occurred, MSDS should be consulted for proper clean up techniques iii) Lifecycle Analysis of Products: - Literature has reported a positive life cycle for the produc tion of biodiesel (1:3). Further research should be done to establish life cycle for this particular   biodiesel process iv) Inherently Safe Design: Goal: to eliminate all hazards in the process using the 10 concepts that follow: 1.

-

Intensification: Use very small amounts of hazardous material so that if there is a leak the hazard will be small. Our plant is of small capacity, and few reactants are of concern

2.

Substitution:  Replace hazardous materials with less hazardous ones. - Our process has eliminated the need to dispose of contaminated water 

3. Attenuation:

8. Status Clear:  It should be possible to see, at a glance, if valves are open or shut, if levels are ok, if correctly assembled. - Process controls will be used on the reactor, mixers, and settling tanks 9. Control: Control systems should be in place. - Process controls will be used on the reactor, mixers, and settling tanks 10. Survival:  If a hazard occurs personnel should be protected. - Determine a fire, explosion, and emergency evacuation plan

Process Safety Management System: -

Employee safety training which includes: o Personal Protective Equipment(PPE) for hearing protection, protective clothing, eye protection, and proper footwear (steel toed CSA approved) o Hazardous materials (eg. METHANOL) used on-site, how and where to locate the MSDS sheets on the materials and on-site via National Fire Protection Agency (NFPA) signs

-

Documentation such as operating manuals, equipment specifications, design criteria will be given to new users to communicate details of plant

-

Employee training in fire, explosion, and emergency evacuation plans

-

Maintenance plans for equipment being installed

Chemical Hazard Information and MSDS:  Definitions:

LD-50/LC-50: Lethal Dose 50. The dose that kills half (50%) of the animals tested. LEL/UEL: Lower/Upper Explosive Limit. Is the limiting/maximum concentration (in air) that is needed for the gas to ignite and explode. TLV: Threshold Limit Value. The reasonable level to which a worker can be exposed without adverse health effects. IDLH: Immediately Dangerous to Life or Health. The exposure to airborne contaminants that is “likely to cause death or immediate or  delayed permanent adverse health effects or prevent escape from such an environment.”

Appendix D Simulation Results

Please Refer to Disc for More Information

Appendix E

Rigorous Calculations For Leaching and Stripping Units

140

141

142

143

144

145