Technoeconomic Analysis of Alternative Pathways of Isopropanol

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TECHNO-ECONOMIC ANALYSIS OF ALTERNATIVE PATHWAYS OF ISOPROPANOL PRODUCTION Warissara Panjapakkul, and Mahmoud M El-Halwagi ACS Sustainable Chem. Eng., Just Accepted Manuscript • DOI: 10.1021/acssuschemeng.8b01606 • Publication Date (Web): 09 Jul 2018 Downloaded from http://pubs.acs.org on July 9, 2018

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ACS Sustainable Chemistry & Engineering

TECHNO-ECONOMIC ANALYSIS OF ALTERNATIVE PATHWAYS OF ISOPROPANOL PRODUCTION

Warissara Panjapakkul1 and Mahmoud M. El-Halwagi1,2

1

2

Chemical Engineering Department, Texas A&M University, College Station, TX, USA, 77843

Gas and Fuels Research Center, Texas A&M Engineering Experiment Station, College Station, TX, USA, 77843

ABSTRACT- Isopropanol is a widely used solvent and chemical. The growing demand for isopropanol, the declining supply of typical feedstocks for manufacturing isopropanol, and the increasing prices necessitate the search for alternative, costeffective, and sustainable pathways for the production of isopropanol. The objective of this work is to synthesize, screen, design, and assess alternate pathways to produce isopropanol. First, biomass- and fossil-based raw materials are considered along with plausible reaction pathways. A process synthesis and integration approach is used to create and prune the alternatives based on branching, matching, and pre-screening. For the promising candidates, process simulation, design, and techno-economic assessment are carried out to compare the options. Under the studied conditions, the results show that propane dehydrogenation followed by direct hydration is the most promising pathway based on profitability (while accounting for price volatility) as well as technical and environmental benefits.

KEYWORDS: Design, process synthesis, process integration, shale gas, glycerin, propane, propylene.

Corresponding Author: M. El-Halwagi. Email: [email protected]

1 ACS Paragon Plus Environment

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INTRODUCTION Isopropanol (isopropyl alcohol or IPA) is an important chemical which is widely used as a solvent, a chemical intermediate, and an ingredient in personal care products, and pharmaceuticals. It is ranked third in lower-alcohol production (following methanol and ethanol)1. The global production of isopropanol is approximately 3*106 metric tons per annum (MTPA), with the US, Europe, and Asia each contributing about 30% of the total production2. The isopropanol market has been steadily growing. Its compound annual growth rate is approximately 7% and the growth is expected to continue especially in the Asia-Pacific region primarily because of the increased use in the pharmaceutical sector3. Isopropanol has been largely produced from propylene via two commercial routes: indirect hydration of refinery-grade propylene and direct hydration of chemical-grade propylene. Indirect hydration is a process with two main reaction steps (esterification and hydrolysis) which convert propylene to isopropanol. The conversion is about 93% and the selectivity to isopropanol and its main byproduct (diisopropyl ether, DIPE) is above 98%4. The main reactions for this process are provided in Table 1. Sulfuric acid is used as a solvent. Therefore, special materials of construction must be used to avoid excessive corrosion. Care should also be given to the treatment and disposal of process wastes including waste water, waste sulfuric acid, spent soda, and off-gases. A key advantage for this process is its ability to use low-purity propylene feed (~40-60 wt%). The indirect-hydration process has been common in the US. Table 1 Chemical reactions of indirect hydration process Step1. Esterification Main reaction

= + ⇆ + = ⇆

Step2. Hydrolysis Main reaction

+ ⇆ +


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+ 2 ⇆ 2 + Side reactions

+ ⇆ + +

⇆ +

The direct hydration process converts propylene to isopropanol via a single-step reaction. It has been the common process used in Europe and Japan2. hindrance primary disadvantage of this method is the requirement for a high-purity feed (at least 90 wt% propylene)4. Furthermore, to enhance the conversion and yield, the reaction is typically operated at much higher temperature and pressure than those used in indirect hydration. On the positive side, direct hydration has some advantages over indirect hydration. The direct hydration process uses water as a solvent instead of a sulfuric acid, thereby reducing the corrosion and environmental problems. The stoichiometric equation of the propylene

hydration

process

is

given

by:

= + ⇆ Similar to the indirect method, the main byproduct of this process is diisopropyl ether (DIPE). There are three types of commercial processes based on the reaction phase of direct hydration.

The recent substantial growth in shale gas production has led to a decline in propylene production from oil refineries.

This decline is attributed to the shift in

ethylene production through ethane dehydrogenation (which produces about 2% of propylene yield) instead of naphtha cracking (which produces about 15% of propylene yield)5-7. The shortfall in naphtha-based propylene along with the rapid growth in propylene-based products caused market shortages and price increase. Therefore, there is



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motivation to investigate various feedstocks and processing pathways for the production of isopropanol. The objective of this work is to synthesize, screen, design, and assess alternate pathways to produce isopropanol from a variety of raw materials. Numerous pathways are first synthesized. High-level techno-economic screening is used to eliminate raw materials and processing pathways that are not economically viable. Next, process simulation, design, and analysis techniques are used to compare the promising options and assess their viability under various conditions.

PROBLEM STATEMENT Isopropanol may be manufactured from various starting materials and via different processing pathways. It is desired to develop a systematic approach for the generation and screening of plausible routes. Because of the need to include sustainability of the supply and viability of the market, fossil- and biomass-based raw materials should be considered. Price volatility is also to be considered in the assessment. Specifically the following questions are to be addressed: •

What are the potential raw materials to be used?

•

What are possible intermediates and byproducts?

•

What are the plausible processing pathways to produce isopropanol?

•

How can the numerous alternatives be quickly and effectively screened?

•

What are the recommended pathways and how susceptible there is their profitability in light of possible price variability?

APPROACH Because of the potentially large number of alternatives for producing isopropanol, a systematic approach is proposed for generating a superstructure of alternative pathways, screening for potential pathways, synthesizing and simulating


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promising flowsheets, performing an economic analysis, and selecting recommended pathways. The following steps summarize the proposed approach: Step i: Generating a superstructure of alternatives The first step is the creation of a superstructure that embeds feedstocks and processing pathways of interest. The superstructure was constructed based on the branching, matching and interception approach introduced by Pham and El-Halwagi8. Branching is a strategy of gathering all of the pathway information by associating the chemicals with sources and main products. The branching approach can be achieved from either a forward branching or a backward branching. The forward branching is a method for searching intermediates from feedstocks by forward approach. For example, in forward branching biomass is a source of methanol, ethanol, etc. Likewise, the backward branching is a method for generating intermediates from products. For example, isopropanol can be made from propylene, acetone, and propane. Therefore, the backward branching should list these chemicals and other species that can be used to make isopropanol. After branching the chemicals, connection of those chemicals or intermediates is needed to form a continuous pathway. When identical chemicals are connected, it is called matching. On the other hand, when different chemicals are connected via chemical reactions, the step is called interception8. Figure 1 is an illustration of the forward and backward branching as well as interception.

.



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Figure 1 An example of forward and backward branching, matching, and interception for sample pathways of isopropanol production

After generating the superstructure of alternatives for isopropanol production, the next step is screening the potential pathways based on a high-level economic analysis. The Metric for Inspecting Sales and Reactants (MISR), an indicator of stoichiometriceconomic “stoichio-nomic” targeting method9 is used for the preliminary screening. MISR9 is defined as:

=

∑

"##$ ∑ !

(1)

where %& is the annual production rate of product ', & is the selling price of product ', %( is the annual feed rate of reactant ) and ( is the purchase cost of reactant ). Pathways with MISR values less than one are eliminated because they are not economically viable (cost of raw materials exceeds value of products). Pathways with MISR values slightly higher than one are not likely to be profitable because of the additional costs associated with capital investment, utilities, labor, and other expenses. When multiple pathways had the same feed and product, a pathway with the least number of steps is selected to promote process simplification.

Step ii: Simulating and designing promising pathways


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For the pathways resulting from the prescreening step, process synthesis, simulation and design studies are carried out to generate the mass and energy balance data for the process, equipment sizing, and key performance indicators. Next, an economic analysis is performed to evaluate feasible pathways. Capital and operating costs are estimate. The return on investment (ROI) is calculated as a measure of profitability. Sustainability and safety issues can also be included in the ROI calculations through the metric of sustainability-and-safety weighted ROI10,11.

Step iii: Selecting the recommended pathway Processes meeting a minimum (threshold) ROI are considered in the final analysis. The margin of error associated with the cost estimation techniques and the impact of price volatility are considered before making a final recommendation on the selected pathway. Figure summarizes the steps involved in the proposed approach.



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Figure 2 Summary of the approach CASE STUDY It is desired to construct a production facility for the production of 375,00 tonnes per year of isopropanol. Various renewable and fossil-based feedstocks as well as processing pathways should be considered. The aforementioned approach is applied as described in the following sections. Generating the superstructure of alternatives Plausible feedstocks, intermediates, byproducts, and wastes were generated along with the processing pathways for conversion. The lists of chemicals and conversion


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technologies are given in Table 2 and Table 3, respectively. In general, numerous feedstocks may be considered based on availability, commercial potential, sustainability, and safety. In this case study, the following fresh feedstocks were considered: acetone, benzene, biomass, butane, ethane, glycerol, methane, methanol, naphtha, and propane. Some of these fresh feedstocks may also be produced as intermediates. Other chemicals were considered to only be intermediates produced within the network. The superstructure of isopropanol production resulting from branching, matching and interception is demonstrated in Figure 3. Table 2 Symbols and prices of chemical species Symbol Chemical Price ($/kg) A Naphtha 0.88 B Biomass 0.30 C Sugar/Carbohydrate 0.57 D Syngas (H2:CO ~ 2) 1.60 E Methane 0.14 F Ethane 0.18 G Glycerol 0.004 H Ethanol 0.78 I Propane 0.48 J Propanol 1.25 K Methanol 0.30 L Acetylene 0.68 M Ethylene 0.39 N Acetaldehyde 1.01 O Butane 0.49 P Benzene 1.21 Q Acetic acid 0.84 R Propylene 0.95 (chemical grade: 94%) purity) 0.75 (refinery grade: 70% purity) S Acetone 1.12 T Isopropyl acetate 1.38 U Isopropanol 1.32



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Table 3 Technology legend of isopropanol production (the reference following each technology designates the source of data and/or flowsheet info) Conversion technology Conversion technology (the number before each technology (the number before each technology represents the pathway arc on represents the pathway arc on Figures 4-6) Figures 4-6) 12 1. Fluid catalytic cracking 20. Fermentation22 2. Fluid catalytic cracking12 21. Syngas fermentation23 13 3. Hydrogenation 22. CO hydrogenation24 4. Transesterification14 23. CO hydrogenation24 14 5. Saponification 24. CO hydrogenation24 6. Hydrolysis14 25. Syngas fermentation23 9 7. Hydrolysis 26. Methanation24 15 8. Pyrolysis 27. CO hydrogenation15 9. Gasification15 28. Stream reforming25 9 10. Landfill 29. Autothermal reforming25 11. Digestion9 30. Dry reforming25 16 12. Pre-hydrolysis 31. Combined reforming25 13. Fermentation17 32. Partial oxidation25 18 14. IB, IBE Fermentation 33. Pyrolysis26 15. Cellobiose degradation19 34. Carbonylation27 20 16. ABE Fermentation 35. Oxidative bromonation28 21 17. Hydrothermal 36. Stream cracking29 18. Fermentation18 37. Stream cracking29 15 19. Gasification 38. Hydro-deoxygenation30 Table 3 Continued Conversion technology (the number before each technology represents the pathway arc on Figures 4-6) 39. Hydro-deoxygenation30 40. Hydrogenolysis31 41. Pyrolysis32 42. Fermentation33 43. Anaerobic digestion33 44. Pyrolysis17 45. Gasification17 46. Stream reforming17 47. Dehydration34 48. Dehydrogenation35 49. Oxidation35 50. Oxidation36

Conversion technology (the number before each technology represents the pathway arc on Figures 4-6) 58. Carbonylation27 59. Hydrogenation39 60. Hydration35 61. Hydroformylation and Hydrogenation4 62. Metathesis12 63. Vapor phase oxidation27 64. Oxidation (Wacker Process)35 65. Oxidation27 66. Oxidation27 67. Cumene and Hock Process20 68. Ketonization40


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51. Direct oxidation37 52. Dehydrogenation38 53. Oxidative dehydrogenation38 54. Dehydration31 55. MTP Process6 56. MTO Process6 57. Hydroformylation35

69. Esterification41 70. Indirect hydration4 (esterification+hydrolysis) 71. Direct hydration4 72. Oxidation20 73. Hydrogenation4 74. Transesterification42 75. Hydrolysis43


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Figure 3 A superstructure of pathways for isopropanol production 12 ACS Paragon Plus Environment

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The process yield was set to the maximum theoretical value9. If a pathway is not economically viable when the yield is maximum, then it can be safely eliminated. For potentially viable pathways, the actual process yield will be used. When the requirement for MISR to be greater than one is used and when an additional constraint on process complexity is added to eliminate any process with more two major technology steps from the starting material to the product (isopropanol), the superstructure shown by Fig. 4 is obtained.

Figure 4 The reduced superstructure after MISR prescreening and process complexity constraint (no more than two major technologies)

Simulating and designing promising pathways Based on the results from the prescreening step, the superstructure may be divided into two categories. The first category includes different routes to produce propylene, which are the glycerol hydro-deoxygenation, propane dehydrogenation (PDH), and methanol to olefins (MTO) process. The second category includes different routes for isopropanol production, which are the indirect hydration of propylene, direct hydration of propylene, and acetone hydrogenation.



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For each promising pathway, simulation was carried out using ASPEN Plus. Whenever actual data for the reaction conversion and yield were available (as will be discussed for the individual pathways), RSTOIC model was used. Otherwise, the reaction was modeled through the minimization of Gibbs free energy (RGIBBS). The results for key pieces of equipment (pumps, compressors, heat exchangers, and distillation columns) were exported to ASPEN Process Economic Analyzer to obtain the equipment cost. The cost of specialized reactors was obtained from literature of similar or analogous units. Lang factors were used to estimate the fixed and total capital investments. The return on investment (ROI) is calculated through the following expression9: ROI =

Annual Net (After - Tax) Profit *100% TCI

(1)

where TCI is the total capital investment of the process and the net profit is given by9: Annual net (after-tax) profit

=

(Annual income – Annual operating cost – Depreciation)*(1-Tax rate) +Depreciation (2a)

= (Annual income – Total annualized cost)*(1-Tax rate) + Depreciation

(2b)

The tax rate was taken as 30% of the taxable income and the depreciation was carried out using a linear scheme over ten years and no salvage value. The process description and the results of process simulation and economic analysis are provided in the ensuing sections. Glycerol hydro-deoxygenation Glycerol hydro-deoxygenation is a one-step catalytic propylene formation reaction with a conversion of 88% and a selectivity of 76%. A key advantage of this reaction is that propylene is the only product in the gas phase so that the separation


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section is not required30. This process is an enticing process because of its abundance and low-price glycerol feedstock. However, the main challenge of this process is that the glycerol from the biodiesel industry is crude and impure, which can damage pipes and equipment44,45. Thus, an expensive glycerol purification technology is required. The process begins with purifying and sending raw glycerol to the reactor to react with hydrogen. The reactor effluent comes out as the vapor-liquid mixture. The mixture is then sent to flash drum to separate propylene, which is the only product in the gas phase, from byproducts, which are all in the liquid phase. The gas stream of propylene and hydrogen is sent to the pressure-swing adsorption section to purify the propylene product and recycle unreacted hydrogen back to the reactor. A flowsheet and stream table of the glycerol hydro-deoxygenation process are presented in Figure 5 and Table A1, respectively. A summary of the key information for economic analysis is demonstrated in Table 4 and the economic results of the glycerol hydro-deoxygenation process are shown in Table 5. Table 4 The key information for the glycerol hydro-deoxygenation process Input/ Output Unit Value Raw glycerol (62 wt%) kt/yr 1,909 Hydrogen kt/yr 0.112 Refinery-grade propylene (75 wt%) kt/yr 361 Fuel (from alcohols) kt/yr 830 Heating utilities MMBtu/hr 652 Cooling utilities MMBtu/hr 847 Electricity kW 638

Table 5 Economic results for the glycerol hydro-deoxygenation process Description Unit Value Fixed capital investment (FCI) MM$ 388 Total capital investment (TCI) MM$ 456 Annual income MM$/yr 349 Annual operating cost MM$/yr 380 -1 ROI yr -2 %



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Figure 5 A flowsheet of glycerol hydro-deoxygenation


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Indirect hydration Refinery-grade propylene gas reacts with sulfuric acid via an esterification reaction to generate a sulfate solution in the absorber reactor. While spent gas is vented out of the process from the top of the absorber reactor, the sulfate mixture from the bottom of the absorber reactor enters the stripper reactor to react with water and form hydration reaction. Unreacted sulfuric acid is removed from the bottom of the stripper reactor and recycled back to the absorber reactor for a further esterification reaction. Product from the hydration reaction is fed to the scrubber to neutralize the residual acid with caustic. Although spent soda is discharged from the bottom of the scrubber, a neutral product leaves the top of the scrubber as vapor stream. The vapor stream is condensed and sent to the drum to separate propylene and propane gases from the product stream. The liquid product stream from the drum is pressurized and fed to the first distillation column to separate DIPE and the remaining gas from the product stream. The distillate is sold as fuel. The product stream then enters the isopropanol-water azeotropic distillation column which uses dimethyl sulfoxide (DMSO) to induce the separation. Isopropanol with a purity of 99 wt% is concentrated in this distillation column and is collected as the product stream. The heavy stream from the isopropanolwater distillation column is sent to the last column to separate DMSO from the aqueous solution to recycle it back to the azeotropic distillation column. Water is discharged from the process as a wastewater. A flowsheet of the indirect hydration process is presented in Figure 6. A stream table of the process is also presented in Table A2. A summary of the key information from the indirect hydration simulation for an economic analysis is demonstrated in Table 6 and the economic results of the indirect hydration process are shown in Table 7.



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Table 6 The key information for the indirect hydration process Input/ Output Unit Value Refinery-grade propylene (70 wt%) kt/yr 452 Stream kt/yr 161 Sulfuric acid aqueous (80 wt%) kt/yr 8.58 Caustic soda kt/yr 96.5 Isopropanol (99.5 wt%) kt/yr 375 Fuel kt/yr 35.1 (from DIPE and light gas) Vent gases with sulfuric acid MMft3/yr 52.6 Process wastewater MMm3/yr 0.06 Spent soda MMm3/yr 0.289 ESP Wastewater MMm3/yr 0.0596 Heating utilities MMBtu/hr 250 Cooling utilities MMBtu/hr 285 Electricity kW 344 Clean water for ESP MMm3/yr 0.0015

Table 7 Economic results for the indirect hydration process Description Unit Value Fixed capital investment (FCI) MM$ 135 Total capital investment (TCI) MM$ 159 Annual income MM$/yr 497 Annual operating cost MM$/yr 422 -1 ROI yr 36 %


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Figure 6 A flowsheet of indirect hydration process



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Direct hydration Liquid chemical-grade propylene enters the reactor and reacts with process water. The liquid effluent is cooled so that propylene and propane vaporize in a flash drum. The vapor stream from the flash drum is condensed and sent to the propylenepropane distillation column to purify unconverted propylene before recycling it back to the feed preparation section. Liquid from the flash drum is sent to the separation section to purify the isopropanol product. The separation section consists of three distillation columns. The first column separates DIPE from isopropanol. Distillate stream from this column is discharged from the process as fuel. The second column is an azeotropic distillation column, which separates water from 99 wt% isopropanol by using DMSO as an extractive solvent. The last column separates water from the solvent in order to recycle water to the feed preparation section and DMSO to the azeotropic distillation column. A flowsheet of the direct hydration process and stream table are presented in Figure 7 and Table A3, respectively. A summary of the key information from the simulation from the direct hydration process for an economic analysis is demonstrated in Table 8 and the economic results of the direct hydration process is shown in Table 9. Table 8 The key information for the direct hydration process Input/ Output Unit Value Chemical-grade Propylene (95 wt%) kt/yr 350 Process water kt/yr 101 Isopropanol (99.8 wt%) kt/yr 326 Fuel (from DIPE and light gas) kt/yr 82 Heating utilities MMBtu/hr 495 Cooling utilities MMBtu/hr 534 Electricity kW 1,253 Table 9 Economic results for the direct hydration process Description Unit Value Fixed capital investment (FCI) MM$ 195 Total capital investment (TCI) MM$ 229 Annual income MM$/yr 432 Annual operating cost MM$/yr 364 -1 ROI yr 23 %


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Figure 7 A flowsheet of direct hydration process 21 ACS Paragon Plus Environment


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Acetone hydrogenation Hydrogen reacts with liquid acetone in the reactor and gives the vapor-liquid phase product. The mixed-phase effluent is sent to drum to separate vapor from the liquid product. The vapor mixture then enters a cooler to condense all the chemicals but hydrogen. This condensed mixture enters the second drum in order to separate unreacted hydrogen from the mixture. The liquid mixture from the second drum combines with the liquid product from the first drum and leaves the process as a 98.5 wt% isopropanol product. Unreacted hydrogen that leaves the second drum is recycled back to the feed preparation section. Figure 8 and Table A4 present a flowsheet and stream summary of the acetone hydrogenation process. A summary of the key information from the acetone hydrogenation simulation for an economic analysis is demonstrated in Table 10 and the economic results of the acetone hydrogenation process is shown in Table 11. It should be noted the price of isopropanol in this process is lower than the isopropanol price in other processes because of its lower purity. The price of isopropanol with 98.5 wt% purity is approximated to be $1.2/kg. Based on the relatively low ROI of 7% (which is less than the minimum acceptable ROI of 10%) and the risk issues associated with the use highpressure hydrogen, acetone hydrogenation was eliminated from further consideration. Table 10 The key information of acetone hydrogenation process Input/ Output Unit Value Acetone (98.7 wt%) kt/yr 359 Hydrogen kt/yr 20.8 Isopropanol (98.5 wt%) kt/yr 371 Fuel kt/yr 11 Heating utilities MMBtu/hr 7.48 Cooling utilities MMBtu/hr 0.525 Electricity kW 884 Table 11 Economic results of acetone hydrogenation process Description Unit Value Fixed capital investment (FCI) MM$ 68 Total capital investment (TCI) MM$ 80 Annual income MM$/yr 444 Annual operating cost MM$/yr 439 -1 ROI yr 7%


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Figure 8 A flowsheet of acetone hydrogenation process



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Selecting the recommended pathway Final selection of propylene production: Since PDH has more potential for profitability compared to MTO12, the pursued pathways of propylene production are PDH and glycerol hydro-deoxygenation. While the economic analysis of glycerol hydro-deoxygenation was estimated from the simulation result, the economic analysis of PDH process was adapted from a recent study7. Based on the data and assumptions in this work, the ROI of the PDH process was calculated to be 20%. (which can slightly increase when sustainability issues are integrated7). Furthermore, the produced hydrogen can be integrated with adjacent facilities to foster industrial symbiosis and to reduce carbon footprint, water uage, and energy utilization. The key information for economic evaluation of PDH process and its results are presented in Table 12 and Table 13, respectively.

Table 12 The key information for the propane dehydrogenation process Input/ Output Unit Value Propane kt/yr 408 Process water kt/yr 101 Propylene (95 wt%) kt/yr 350 Hydrogen kt/yr 30 Heating utilities MMBtu/hr 1220 Cooling utilities MMBtu/hr 733 Electricity kW 21372

Table 13 Economic results for the propane dehydrogenation process Description Unit Value Fixed capital investment (FCI) MM$ 430 Total capital investment (TCI) MM$ 507 Annual income MM$/yr 363 Annual operating cost MM$/yr 234 Annual ROI yr-1 20 %


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Based on the economic analysis of the glycerol hydro-deoxygenation, the annual ROI of the process was negative (-2%), which indicates that it is not economically viable under the studies conditions. Additionally, there are some issues of glycerol hydro-deoxygenation process such as insufficient amount of glycerol from biodiesel production, and uncertainty in process scale-up. Therefore, it can be concluded that PDH is the best pathway for propylene production. Although the ROI result of the glycerol hydro-deoxygenation using chemical market prices is negative, the ROI of this process can be positive if the price of raw glycerol decreases and the price of propylene increases. From the sensitivity analysis of hydro-deoxygenation from raw glycerol in Figure 9a, the process can be operate economically with raw glycerol prices ranging from $0.06/kg to $0.13/kg when the propylene price is high (more than $1.1/kg). In order to compare the profitability of glycerol hydro-deoxygenation from raw glycerol and from refined glycerol, the sensitivity analysis of the process from refined glycerol was also performed. According to the sensitivity analysis shown in Figure 9b, the range of refined glycerol price for economically feasible scenario is between $0.32-0.45/kg, depending on the propylene selling price. However, it should be realized that the market price of purified glycerol is $0.85/kg and the market price of propylene is in the range of $0.9-1.2/kg. This implies that propylene production from refined glycerol is hardly economically viable because the market price of refined glycerol is about twice the highest feasible price of refined glycerol for the propylene market price scenario. Therefore, it can be concluded that the using raw glycerol is preferred to using pure glycerol to produce propylene via hydro-deoxygenation. Nonetheless,

50 40 30 $1.1/kg 20

$1.3/kg

10

$1.5/kg

0 0.05

0.1

0.15

0.2

0.25

Price of raw glycerol ($/kg)

Annual ROI (percent)

since the base case has a negative ROI, this option is not pursued anymore. Annual ROI (percent)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


50 40 30 $1.1/kg 20

$1.3/kg

10

$1.5/kg

0 0.3

0.35

0.4

0.45

0.5

Price of purified glycerol ($/kg)

(a)

(b) 25 ACS Paragon Plus Environment


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Figure 9 Sensitivity analysis for the annual ROI of hydro-deoxygenation from raw glycerol (a) and from purified glycerol (b)

Final selection of isopropanol production: While the ROI of the indirect hydration was equal to 36%, the ROI of the direct hydration was equal to 23%. As the margin of error in in these types of conceptual design and economic studies is typically in the order of 20-25%, it cannot be fully concluded which pathway is the best pathway to produce isopropanol. Additionally, other sustainability and safety issues should be considered. A sensitivity analysis was performed to assess the impact of varying propylene and isopropanol prices on ROI. The results of the sensitivity analysis for the indirect and the direct hydration processes are shown in Figure 10. The results from the sensitivity analysis show that indirect hydration and direct hydration processes are economically feasible with the isopropanol market price of $1.323/kg when the price of the refinery–grade propylene drops below $0.93/kg and the price of the chemical-grade propylene drops below $1.17/kg, respectively. Additionally, when the price of the refinery-grade propylene is reduced to a range of $0.61-1.03/kg, the minimum acceptable level of ROI (10%) for the indirect hydration process is achieved. Similarly, the minimum acceptable level of ROI of the direct hydration process is reached when the price of the chemical-grade propylene is decreased to a range of $0.77-1.24/kg for the different product selling prices. Comparing the impacts of the propylene price to the ROI of indirect hydration and direct hydration processes, the ROI of the direct hydration process is less dependent on its raw material price than the ROI of the indirect hydration process. Hence, the indirect hydration process is much more sensitive to price volatility and the ROI may easily drop below the minimum acceptable with plausible variations in prices. The direct hydration process shows a lower slope of sensitivity and is, therefore, more economically robust under price-changing conditions. Furthermore, the indirect 26 ACS Paragon Plus Environment

Page 27 of 41

hydration process discharges wastewater stream that are laden with sulfuric acid and caustic soda. As such, from an environmental sustainability perspective, direct hydration offers key advantages. Therefore, while both routes are recommended, the direct route is favored given the

70 60 50 40 30 20 10 0

$1/kg $1.323/kg $1.5/kg 0.5 0.6 0.7 0.8 0.9 1 1.1 1.2

Price of refinery-grade propylene ($/kg)


sensitivity and sustainability issues. Annual ROI (percent)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


70 60 50 40 30 20 10 0

$1/kg $1.323/kg $1.5/kg 0.7 0.8 0.9 1 1.1 1.2 1.3 1.4

Price of chemical-grade propylene ($/kg)

(a)

(b)

Figure 10 Sensitivity analysis for the annual ROI of indirect hydration (a) and direct hydration (b) processes (the lines represent the ROI for various selling prices of isopropanol)

Since propylene hydration is recommended as the primary route for isopropanol production, the options of making versus buying propylene were considered. The ROI for the “buy” scenario followed by direct hydration process is 23%. The ROI for the “make” case, which was calculated based on the economic analysis of direct hydration and PDH processes, was found to be 22%. A summary of economic results of purchasing and manufacturing scenarios is presented in Table 14. Table 14 A summary of economic results of two possible scenarios Propylene “make” Description Propylene scenario “buy” scenario Fixed capital investment (MM$/yr) 195 625 Total capital investment (MM$/yr) 229 736 Annual income (MM$/yr) 432 462 Annual operating cost (MM$/yr) 364 262 ROI (yr-1) 23 % 22 % 27 ACS Paragon Plus Environment


Due to the very close ROI values for the two scenarios, a sensitivity analysis was carried out, as shown in Figure 11, to assess the changes resulting from variations in propylene (in the buy scenario) and propane (in the make scenario) prices. The results illustrate that the ROI of the buy scenario is more sensitive to the changes in propylene price than the ROI of the make scenario to the propane price. Thus, it was concluded that it is more economically robust to produce make rather than buy propylene. The sensitivity analysis of the make scenario also reveals that such route can operate economically when the propane price is less than $0.95/kg at the isopropanol price of $1.23/kg. Additionally, the minimum acceptable level of ROI of 10% for the make case is achieved with the propane price between $0.45-0.83/kg, depending on the isopropanol selling price. It is also worth noting that when sustainability issues (including water usage, fuel consumption, carbon footprint, and volatile organic compound emission) are included, the sustainability weighted return on investment10 is enhanced by 1 to 6 yr-1% depending on the type of process modification undertaken to integrate and intensify the process7. Further improvements can also be made when the process is integrated with adjacent facilities

70 60 50 40 30 20 10 0

$1/kg $1.323/kg $1.5/kg 0.7

0.8 0.9

1

1.1

1.2


through an industrial symbiosis framework46-49.


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 28 of 41

70 60 50 40

$1/kg

30 20

$1.323/kg

10

$1.5/kg

0 0.3 0.4 0.5 0.6 0.7 0.8

Price of chemical-grade propylene ($/kg)

Price of propane ($/kg)

(a)

(b)

Figure 11 Sensitivity analysis for the annual ROI of (a) propylene “make” scenario and (b) propylene “buy” scenario

CONCLUSIONS


Page 29 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


This work focused on the production of isopropanol from various renewable and fossilbased feedstocks. A superstructure was created to embed a rich set of potential pathways. The superstructure was constructed based on the branching, matching and interception approach. Prescreening using high-level techno-economic criteria was carried to eliminate infeasible or non-competitive pathways. The remaining promising pathways included: propylene direct and indirect hydration (where propylene may be produced through glycerol hydro-deoxygenation, propane dehydrogenation, or methanol to olefins process) and acetone hydrogenation. The results of flowsheet synthesis, process simulation, techno-economic analysis, and other sustainability considerations were used to compare the alternate pathways. Under the studied conditions, it was found that propylene production via propane dehydrogenation is superior to glycerol hydro-deoxygenation and the most profitable pathway to produce isopropanol is direct hydration of propylene. A sensitivity analysis was carried out to confirm the conclusions under potential scenarios of price volatility.

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Page 34 of 41

APPENDIX Table A1 A stream table of glycerol hydro-deoxygenation Stream Name

Units

1

2

Temperature

K

323

559

Pressure

bar

20

Mass Flows Mass Fractions Hydrogen

kg/hr

7

3

4

5

6

298

573

298

477

80

80

80

1

7

161218

161218

238639

1.000

1.000

1.000

1.000

Propylene

0.000

0.000

0.000

Methanol

0.000

0.000

0.000

2-Propenol

0.000

0.000

1-Propanol

0.000

2-Propanol Water Propylene glycol 1,3 Propanediol

7

8

9

10

479

573

573

1

80

80

80

80

80

80

143723

143723

143723

304941

304941

201149

103792

0.000

0.000

0.000

0.000

0.512

0.512

0.776

0.000

0.000

0.000

0.000

0.000

0.000

0.144

0.144

0.217

0.002

0.000

0.335

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.008

0.008

0.001

0.022

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.022

0.022

0.002

0.060

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.008

0.008

0.001

0.021

0.000

0.000

0.000

0.000

0.000

0.003

0.003

0.003

0.217

0.217

0.004

0.629

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.017

0.017

0.000

0.050

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.017

0.017

0.000

0.050

298

11 298

12 298

Glycerol

0.000

0.000

0.000

0.000

0.617

0.997

0.997

0.997

0.056

0.056

0.000

0.166

NaOCH3

0.000

0.000

0.000

0.000

0.027

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Triglycerides

0.000

0.000

0.000

0.000

0.020

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Protein

0.000

0.000

0.000

0.000

0.001

0.000

0.000

0.000

0.000

0.000

0.000

0.000


Page 35 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47


Table A1 Continued Stream Name

Units

13

14

Temperature

K

298

298

Pressure

bar

80

80

Mass Flows Mass Fractions Hydrogen

kg/hr

161211

45178.2

1.000

0.000

Propylene

0.000

0.965

Methanol

0.000

0.000

2-Propenol

0.000

0.003

1-Propanol

0.000

0.008

2-Propanol

0.000

0.006

Water Propylene glycol

0.000

0.018

0.000

0.000

1,3 Propanediol

0.000

0.000

Glycerol

0.000

0.000

NaOCH3

0.000

0.000

Triglycerides

0.000

0.000

Protein

0.000

0.000



Page 36 of 41

Table A2 A stream table of indirect hydration process Stream Name

Units

1

2

3

4

5

6

7

8

9

10

11

Temperature

K

363

363

363

372

367

342

407

407

406

422

363

Pressure

bar

6

6

6

6

6

1.01325

1.01325

3

1.01325

1.01325

1.01325

Mass Flows

kg/hr

1072

1072

56508.3

1672.86

89697.4

89697.4

89697.4

20170.1

87556.8

63956.3

63956.3

Mass Fractions Propylene

0.000

0.000

0.700

0.210

0.027

0.027

0.027

0.000

0.027

0.000

0.000

Propane

0.000

0.000

0.300

0.664

0.177

0.177

0.177

0.000

0.181

0.000

0.000

DIPE

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.014

0.000

0.000

Isopropanol Diisopropyl sulfate

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.577

0.003

0.003

0.000

0.000

0.000

0.094

0.331

0.331

0.331

0.000

0.000

0.000

0.000

Water Isopropyl hydrogen sulfate

0.200

0.200

0.000

0.032

0.171

0.171

0.171

1.000

0.196

0.048

0.048

0.000

0.000

0.000

0.001

0.248

0.248

0.248

0.000

0.006

0.000

0.000

DMSO

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Sulfuric acid

0.800

0.800

0.000

0.000

0.047

0.047

0.047

0.000

0.000

0.949

0.949

Caustic soda

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Sodium bisulfate

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000


Page 37 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47



Units

12

13

14

15

16

17

18

19

20

21

22

Temperature

K

363

407

418

407

325

325

325

326

312

312

Pressure

bar

6

1.01325

1.01325

1.01325

1.01325

1.01325

1.01325

5.06625

4.053

4.053

4.32879

Mass Flows

kg/hr

63956.3

12064

85538.4

14531.7

85550.5

18172.1

67378.4

67378.4

3696.35

697.383

62984.6

0.000

0.000

0.027

0.001

0.027

0.106

0.006

0.006

0.100

0.024

0.000


395

Propane

0.000

0.000

0.184

0.005

0.184

0.698

0.045

0.045

0.782

0.228

0.000

DIPE

0.000

0.000

0.014

0.002

0.014

0.008

0.016

0.016

0.113

0.668

0.003


0.003

0.000

0.582

0.080

0.582

0.159

0.696

0.696

0.002

0.032

0.745

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000


0.048

0.310

0.192

0.048

0.192

0.028

0.237

0.237

0.003

0.048

0.252

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

DMSO

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Sulfuric acid

0.949

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Caustic soda

0.000

0.690

0.000

0.066

0.000

0.000

0.000

0.000

0.000

0.000

0.000

Sodium bisulfate

0.000

0.000

0.000

0.798

0.000

0.000

0.000

0.000

0.000

0.000

0.000



Page 38 of 41


Units

23

24

25

26

27

Temperature

K

345

355

424

371

474

Pressure

bar

1.31

1.03421

1.31

1.03421

1.31

Mass Flows

kg/hr

132838

46923.2

148891

16053.1

132838

0.000

0.000

0.000

0.000

0.000

Mass Fractions Propylene Propane

0.000

0.000

0.000

0.000

0.000

DIPE

0.000

0.004

0.000

0.000

0.000


0.000

0.995

0.001

0.011

0.000

0.000

0.000

0.000

0.000

0.000


0.000

0.000

0.107

0.989

0.000

0.000

0.000

0.000

0.000

0.000

DMSO

1.000

0.000

0.892

0.000

1.000

Sulfuric acid

0.000

0.000

0.000

0.000

0.000

Caustic soda

0.000

0.000

0.000

0.000

0.000

Sodium bisulfate

0.000

0.000

0.000

0.000

0.000


Page 39 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47


Table A3 A stream table of direct hydration process Stream Name

Units

1

2

3

4

5

6

7

8

9

10

11

Temperature

K

316

323

395

308

308

311

543

508

437

437

437

Pressure

bar

17.51

17.51

151.99

1.01

1.01

151.99

151.99

151.99

151.99

17.37

17.37

Mass Flows

kg/hr

43703.8

54918.3

54918.3

12610.7

50010.4

50010.4

50010.4

104928

104928

16374.8

88553.7


0.948

0.754

0.754

0.000

0.000

0.000

0.000

0.118

0.118

0.280

0.089

Propane

0.052

0.046

0.046

0.000

0.000

0.000

0.000

0.024

0.024

0.051

0.019

DIPE

0.000

0.047

0.047

1.000

1.000

1.000

1.000

0.384

0.384

0.158

0.425

Isopropanol

0.000

0.148

0.148

0.000

0.000

0.000

0.000

0.468

0.468

0.495

0.463

Water

0.000

0.005

0.005

0.000

0.000

0.000

0.000

0.006

0.006

0.015

0.004

DMSO

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

12

13

14

15

16

17

18

20

21

22

Stream Name

Units

356

416

374

474

23

Temperature

K

328

316

446

316

316

403

Pressure

bar

17.37

17.24

17.51

5.07

5.07

5.34

1.31

1.03

1.31

1.03

1.31

1.01

Mass Flows

kg/hr

16374.8

5160.17

11214.6

10020.5

226.813

78306.3

221015

40779.4

258415

37399.7

221015

37399.7

0.280

0.888

0.000

0.777

0.244

0.000

0.000

0.000

0.000

0.000

0.000

0.000


345

19

308

Propane

0.051

0.112

0.023

0.169

0.064

0.000

0.000

0.000

0.000

0.000

0.000

0.000

DIPE

0.158

0.000

0.231

0.003

0.079

0.480

0.000

0.002

0.145

1.000

0.001

1.000

Isopropanol

0.495

0.000

0.723

0.020

0.462

0.520

0.000

0.998

0.000

0.000

0.000

0.000

Water

0.015

0.000

0.022

0.032

0.151

0.000

0.000

0.000

0.000

0.000

0.000

0.000

DMSO

0.000

0.000

0.000

0.000

0.000

0.000

1.000

0.000

0.855

0.000

0.999

0.000



Page 40 of 41

Table A4 A stream table of acetone hydrogenation process Stream Name

Units

1

2

3

4

5

6

7

8

Temperature

K

323

380

365

373

293

295

373

373

Pressure

bar

20

30

30

30

1.01325

30

30

30

Mass Flows

kg/hr

2623.95

2623.95

3352.72

3352.72

44876.2

44876.2

44876.2

48228.9

Mass Fractions Hydrogen

1.000

1.000

1.000

1.000

0.000

0.000

0.000

0.038

Acetone

0.000

0.000

0.000

0.000

0.987

0.987

0.987

0.001

Isopropanol

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.948

Water

0.000

0.000

0.000

0.000

0.013

0.013

0.013

0.012

2-Hexanol

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.001

10

11

12

13

14

15

16

Stream Name

Units

9

Temperature

K

373

373

313

313

313

313

313

368

Pressure

bar

30

30

30

30

30

30

30

30

Mass Flows

kg/hr

6223.04

42005.9

6223.04

2103.13

4119.91

728.773

1374.35

46125.8

Mass Fractions HYDRO-01

0.289

0.000

0.289

0.855

0.000

1.000

0.77851

0.000

ACETO-01

0.002

0.001

0.002

0.002

0.002

0.000

0.00237

0.001

ISOPR-01

0.697

0.986

0.697

0.140

0.981

0.000

0.21405

0.985

WATER

0.012

0.012

0.012

0.003

0.017

0.000

0.00507

0.012

HEXYL-01

0.000

0.001

0.000

0.000

0.000

0.000

5.5E-09

0.001


Page 41 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47


For Table of Contents Use Only

TOC Graphic

Synopsis: This figure is a superstructure that embeds various pathways for producing isopropanol from several feedstocks and intermediates.


Technoeconomic Analysis of Alternative Pathways of Isopropanol

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