The economics of the process is strongly dependent on the price of acetic acid, and we examined scenarios under which acetic acid was available at either $0.16/lb or $0.12/lb. The total capital investment in either situation is approximately $47,000,000. If acetic acid is available at $0. 16/1b, we estimate an IRR of 11.3 %, but if acetic acid can be purchased for $0.12/Ib the IRR is 18.5% after 20 years. It is our recommendation to pursue more research into projecting both the cost of acetic acid and the market for acetaldehyde. If acetic acid will be available at the lower price, the company should pursue production of acetaldehyde.

The process design consists of three major components: pre-treatment, reaction and separation. In the pre-treatment stage, the switchgrass is ground and cellulose is extracted using highly dilute base. The reaction occurs in water solvent under 5 MPa of hydrogen in a fixed bed catalytic reactor with a retention time of 30 minutes. The ethylene glycol product is separated from the solvent and byproducts by flash vessels and column distillation in order to attain 99.5 % purity. A side product, propylene glycol, is also produced at 99.5 % purity and is separated in a similar fashion.

For the economic analysis, the plant was assumed to be located in Tennessee (near the Gulf Coast) where switchgrass is readily available. The total capital investment is $349,000,000, including a working capital of $7,747,000. The process is sensitive to the amount of solvent used. In the base case scenario using the same amount of solvent as used in the patent with ethylene glycol priced at $0.50/lb, propylene glycol byproduct priced at $0.78/lb, switchgrass priced at $50/dry ton, hydrogen priced at $1.00/lb, and catalyst priced at $5.08/lb, the net present value (NPV) of the project is negative $565,500,000 based on an interest rate of 15%, and the investor’s rate of return (IRR) is negative. The process is not profitable under the current conditions but may become profitable with advances in technology as described by the sensitivity analysis section of the report. Although the current version of the process is unprofitable, it is expect that the further development of the technology described in the referenced patent will allow for profitable future versions of this process, and it is recommended by the design team that the management retain this report for future reference

2600 tons per day of coal and 66 million standard cubic feet per day of natural gas are converted to syngas in separate, parallel process trains. The hydrogen rich natural gas syngas is mixed with the hydrogen lean coal syngas to give the desired syngas composition. Fischer-Tropsch chemistry is used to convert syngas with a 2 to 1 H₂:CO molar ratio to a distribution of alkanes. The alkanes are separated to give 15,500 barrels per day of liquid product.

The total capital investment of the project was estimated to be $1,308 million. Using the EIA baseline projection for the price of oil and a 15% discount rate, the project is expected to have a net present value of -$258 million, with an internal rate of return of 11%. The profitability of the project is especially dependent on the price of oil and the total capital investment.

This report provides a design and economic analysis for the production of butadiene on the Gulf Coast. Process flow sheets, energy and utility requirements, and equipment summaries are provided and analyzed. Process profitability is sensitive to the cost of both ethanol and butadiene. It is shown that the plant is very profitable for its expected 15-year lifespan with an expected internal rate of return of 40%, return on investment of 34%, and net present value of $172,000,000 (for a discount rate of 15%). The process becomes unprofitable if the price of ethanol increases to over $3.00/gallon. A combination of increased ethanol price and decreased butadiene price will also cause the plant to be unprofitable. Therefore, plant construction is only recommended given an acceptable price of ethanol and butadiene.

While the existing platform successfully provides power to the overall liquefaction process, the gaseous exhaust from this process leaves the system at elevated temperatures. The processes presented in this project seek to recover the heat that is lost through the exhaust and therefore, improve the thermodynamic efficiency of this system. Additionally, these processes more rigorously meet environmental standards concerning flue gas compositions and temperatures. Seven such processes are presented in this report. Each of these provides a net of 40MW, the required power to drive the liquefaction process, while performing at higher thermodynamic efficiencies than the simple gas turbine process.

Rigorous economic analyses were performed for each of the presented processes. One recovery system has a lower net present value (NPV) than that of the simple gas turbine, four have approximately equal NPVs, and two systems have significantly better NPVs than that of the simple gas turbine. The optimal system has an NPV of $22 million and an internal rate of return (IRR) of 28.2% versus the simple gas turbine with an NPV of $12.3 million and an IRR of 20.3%. Further analyses of the economic and pricing assumptions may be required before final project approval.

A major challenge of the project is the non-ideal behavior of the components, which includes multiple azeotropes and distillation boundaries. Another important characteristic of this design project is the unusually small scale: one 6,240 gallon quantity of used co-solvent mixture must be processed every two days. Due to this scale, batch processes were investigated as well as continuous processes.

The continuous alternative utilizes three major separation units: a 15-tray distillation column, a decanter for the distillate, and an evaporator for the bottoms. 93.8% of the original IPA and 99.2% of the original NPB is recovered. There is no oleic acid and 0.5% by mole of water in the product. Pure NPB and IPA are added at the end of the separation to compensate for the lost co-solvents, and to restore the IPA/NPB ratio to 56/44. A $0.45/lb selling price of reclaimed co-solvent returns an IRR of 45.6% and an NPV of $2,657,300 in year 10 of production.

The batch alternative utilizes a batch distillation column with multiple receivers and recovers 95.4% of the original IPA and 99.7% of the original NPB. There is 0.8% water by mole and no oleic acid in the product. Pure NPB and IPA also must be added to restore the original ratio. Because the composition of water is higher in the batch product than in the continuous, the co-solvent mixture is sold at a lower $0.42/lb, resulting in an IRR of 37.2% and an NPV of $2,452,500 in year 10. However, the batch process has significant down time and can potentially handle up to four times the solvent demand (2 trucks of solvent/day), resulting in an IRR of 130% and an NPV of $22,494,500 at the same selling price. Because of the batch plant’s ability to handle demand growth, its flexibility in separating different co-solvent ratios, and its robust economic potential, we recommend the construction of the batch co-solvent reclamation plant.

This design analyzes two proposed locations, the US Midwest or Brazil, and their associated renewable feedstocks, corn or sugarcane juice, respectively. This report investigates the relative economic attractiveness of each option. The US case requires location near an existing industrial ethanol fermentation plant to give easy access to dry-ground corn as a carbohydrate source. This case yields an IRR of 17.56% and an overall NPV of $35.2 million at a 15% discount rate. The Brazil case has comparatively cheaper feedstock, however because of seasonality and total usable carbohydrate content, it requires a greater mass of feedstock and increased capital investment relative to the US case. The NPV difference of the two cases is extremely sensitive to the assumed price of sugarcane juice which has recently shown extraordinary volatility. Based on this analysis, the US location seems most promising; however, detailed laboratory level studies are needed to confirm the profitability and assumptions made.

The process is designed for SCM and consumes a total feed of 9.35 billion pounds of molasses per year. Corn dry grind is simply too expensive, and biomass, while cheaper per pound, imposes too many additional pre-processing costs. The molasses first undergoes hydrolysis then hydrogenation, followed by condensation and separation involving distillation and crystallization. Transalkylation and aqueous phase reforming are also employed to boost yield and create a self-contained process.

Several key assumptions are inherent in this process’s design. First, all reactor yields come directly from specific examples in the literature. Second, results found in the patents for glycerol were assumed valid for sorbitol as well, since not all patents used the same materials for their examples. Third, the economic analysis assumes that raw materials for catalyst manufacture can be purchased in bulk for a quarter of the price for small quantities. This assumption was suggested by Dr. Fabiano.

Based on these assumptions, the process designed herein meets the desired non-financial criteria, but results in an investor’s rate of return of negative 2.90% and a net present value of negative $196 million. However, further research into the catalyst or reactor yields could easily allow the process to break even or offer an attractive return.

This product, known as the Diamond HemeSep blood processing unit, could offer more standardized, efficient blood separation technologies that would benefit health care providers, patients and researchers. Moreover, the product is predicted to have a healthy financial outlook: based on the target market of clinical laboratories performing preclinical and clinical trials involving numerous samples of blood, we expect to sell 1 million cartridges in the first year of production with sales growing to 1.7 million cartridges in the tenth and final year. The net present value (NPV) of the proposed project, based on a selling price of $25 a cartridge, is expected to be $51 million. For the current projections, Series A investors can expect returns of 45%.

This report provides a design and economic analysis for the production of butadiene on the Gulf Coast. Process flow sheets, energy and utility requirements, and equipment summaries are provided and analyzed. Process profitability is sensitive to the cost of both ethanol and butadiene. It is shown that the plant is very profitable for its expected 15-year lifespan with an expected internal rate of return of 40%, return on investment of 34%, and net present value of $172,000,000 (for a discount rate of 15%). The process becomes unprofitable if the price of ethanol increases to over $3.00/gallon. A combination of increased ethanol price and decreased butadiene price will also cause the plant to be unprofitable. Therefore, plant construction is only recommended given an acceptable price of ethanol and butadiene.

A separations process incorporating a membrane presents a unique opportunity for our company to penetrate the market for olefin separations. Possible membrane separation processes include a membrane-only separation process and a hybrid system, the latter consisting of both membranes and distillation columns. The costs and benefits of each process were analyzed to determine the price of the membrane that would make a membrane separation process competitive. More specifically, the conventional distillation processes were examined as base cases to establish the target internal rate of return. Then the price of membrane that would yield the same IRR as the conventional distillation process was calculated for both membrane only and hybrid processes in each of the three cases below. For each of the three cases below, the conventional distillation separation process was established as the base case while the membrane only separation process and the hybrid system were explored as alternatives.

The cases below represent the feed and product combinations desired for the various separation processes that were explored.

• Case A: Chemical grade (93%) to polymer grade (99.5%) propylene, with the chemical grade propylene feed coming from a steam cracker
• Case B: Refinery grade (35%) to polymer grade (99.5%) propylene, with the refinery grade propylene feed coming from the propane dehydrogenation process
• Case C: Refinery grade (35%) to chemical grade (93.0%) propylene, with the refinery grade propylene feed coming from the propane dehydrogenation process

After examining the three cases, the hybrid system was not considered a viable option for separation in case B and case C. The price of the membrane is obtained when the IRR of the membrane only and/or hybrid process are equal to that of conventional distillation. The price of the membrane was determined to be in the price range of $47-52/m². This price was capped at the lowest membrane price calculated for any process ($52), for reasons explained in detail in the Conclusion. However, the membrane product must be priced lower than $52, because the IRR of a membrane process must be higher than that of conventional distillation in order to give our clients a compelling enough reason to adopt our membrane technology.

The algae-to-biofuel venture was segmented into three modules: algal cultivation, lipid extraction and lipid processing. Each module was studied thoroughly and several strategies were proposed for the reduction of its associated fixed, capital and variable costs. As contrasted with a previous study, it was concluded that heterotrophic algal cultivation and transesterification lipid- processing technologies would improve the efficiency and reduce the total capital investment. Once each module was designed in detail, the three segments were stitched together to perform an overall economic analysis. Based on the current market price of $3.30 per gallon for pure biodiesel, a project life of 15 years, and a 15% discount rate, the results indicate that an algae-to- biodiesel process may not only be profitable, but also a sound and reasonable investment. The project’s projected Net Present Value (NPV) is $1.3 billion and the Return on Investment (ROI) was determined to be 32%.

Although these economic results are promising, they are based on an analysis that necessarily invoked highly uncertain postulates in the dearth of published data. For example, the kinetics used to model the lipid-processing module were based on data collected for palm oil at similar conditions, while the details of lipid-extraction energy usage for a high-density slurry were approximated on the basis of results for low-density slurry. Furthermore, it was concluded that the income from the sale of the algal biomass byproduct of lipid-extraction is a critical factor in the profitability. Based on its protein content, this report considered the use of algal biomass as animal feed and determined its economic worth accordingly. However, to ensure the economic success of biodiesel production, an additional analysis should focus on algal usage of biomass as a feedstock and confirm the safety of its use. Further analyses could examine other potential applications for the byproduct, including opportunities within the pharmaceutical and power generation industries. Overall, in order to convince investors that the attractive economics published in this report may be translated into actual earnings, it is critical to move beyond modeling. Pilot studies must be conducted in order to bolster the proposed algae-to-biodiesel venture with experimental data and identify possible pitfalls.

To obtain useful information about the feasibility and scalability of the process, 30 M kg/yr of each product will be produced. The products will be tested for purity and samples will be sent out to consumers to demonstrate the quality of the product. The MAA and MEK must be of the same purity generated by current commercial processes. The pilot plant will be designed in three major parts. The first part consists of the bacterial fermentors that are used to produce and scale up MAA and MEK production. Relatively little is currently known about the efficiency of production of MAA and MEK by E. coli and this part of the plant will provide critical data about conditions required for the bacteria as well as production rates. The second part of the plant consists of the MAA purification process. Many options will be considered for the purification steps, many of which will have to be modeled in ASPEN because MAA is usually not produced in the aqueous phase. The final section of the plant will be used for MEK purification. To reduce plant costs, the design will try to share equipment between the two purification processes.

The main goal of the plant is to obtain data and demonstrate feasibility, not to demonstrate sustainable profitability. Estimates for total capital investment and show that the plant will not be profitable for the first five years of operation, but the valuable data gained from the operation will be used to design the larger, more efficient, full-scale plant. The total capital investment required for the plant is approximately $ 6.33 million. Returns generated from sales are minimal compared to the capital investment and operating costs. A full scale plant is expected to be profitable over time because of economies of scale and the price of inputs and outputs of the process.

The current water supply system in Las Delicias consists of 75 hp (56.25 kW) pump system. It was determined that a re-design of the hydraulic system was necessary to reduce power requirements. The new design added a new holding tank and eliminated the need for a 60 hp booster pump, reducing the total pump power needs to 34 hp (25.5 kW). This design allows the villagers to receive continuous water. The total investment of this new design is $120,000 and yields a NPV of $413,000 (at 1.6% discount rate) and an IRR of 36%.

The system in Apatut is a grass roots project. We worked with the initial design provided by the EWB-MAP team that is currently involved with the project. The total investment is $22,000 and yields a NPV of $78,000 (at 1.6% discount rate) and an IRR of 41%.

In the electrolysis of steam to form hydrogen, the hydrogen is stored in pressurized vessels at 100 psi. During fuel cell mode, the hydrogen gas is oxidized in the fuel cell to form water, which is again stored for use in electrolysis the following day. The operating temperature for both electrolysis and fuel cell operation is 1472°F. The overall efficiency for the hydrogen system is 52.4%, with the main losses occurring as heat supplied to the electrolyzer.

The second design electrolyzes a mixture of steam to hydrogen and carbon dioxide to carbon monoxide at 1292°F and 147 psia. The resulting syngas is fed through a methanation reactor to produce a methane rich stream. The overall efficiency of the methane-based system is 55.7%, with the main losses coming from compression and heating for electrolysis mode.

The molten antimony fuel cell uses an equimolar mixture of antimony and antimony trioxide as the feedstock for electrolysis. The electric current in the electrolyzer reduces the antimony trioxide to form a stream of pure molten antimony. Both the electrolyzer and fuel cell operate isothermally at 1292°F and 14.7 psia. The overall efficiency for the antimony design is 53.7%, with the main losses coming from heat supplied to the electrolyzer.

For profitability analysis, off peak electricity was priced at $0.06/kWh and peak power was priced at $0.20/kWh. Under these optimistic assumptions, the return on investment (ROI) for the hydrogen design was calculated to be -26.1%. For the methane system, the ROI was calculated to be -19.2%. For the antimony case, the ROI was found to be -34.2%. These designs serve as a framework for future work with electrical energy storage. However, we believe that with improvements in system efficiency and reductions in the initial capital investment, future reversible fuel cell systems will be profitable and competitive with other forms of electrical energy storage.

The economic analysis that follows assumes a grassroots plant on the US Gulf Coast. The total capital investment was calculated to be $34.0 million, which includes a working capital of $9.78 million. Under the assumptions that the prices of crude glycerol, hydrogen, and propylene glycol are $0.22, $0.50, and $1.00 per pound respectively, the net present value (NPV) at the end of the 15 year allotted course of the project is $88.4 million and the investors’ rate of return (IRR) is 58.45%. The price of glycerol is projected to remain stable or decrease in the future and the price of propylene glycol is projected to remain stable or increase suggesting that this project could become even more profitable in the future.

The apparent profitability of this project is largely caused by the efficient and cost effective method of desalting glycerol through electrodeionization. Unfortunately, the proprietary nature of this new process precludes public access to true costing and specifications for the equipment since firms as Dow largely control the technology. Thus, conservative estimations were made in our economic analysis to account for this uncertainty.