Techno-economic assessment of an energy self-sufficient process to produce biodiesel under supercritical conditions

Depósito de Investigación de la Universidad de Sevilla https://idus.us.es/ This is an Accepted Manuscript of an article published by Elsevier in The Journal of Supercritical Fluids, Vol. 128, on October 2017, available at: https://doi.org/10.1016/j.supflu.2017.03.010 © 2017 Elsevier. En idUS Licencia Creative Commons CC BY-NC-ND 1 Techno-economic assessment of an energy self-sufficient process to produce biodiesel under supercritical conditions F. J. Gutiérrez Ortiz, P. de Santa-Ana Departamento de Ingeniería Química y Ambiental, ETS de Ingeniería, Universidad de Sevilla Camino de los Descubrimientos, s/n. 41092 Sevilla, Spain Phone no: + 34 95 448 72 68 / 60 e-mail: “Francisco Javier Gutiérrez Ortiz” fra[email protected] Abstract This paper presents a conceptual design, simulation and techno-economic assessment of an energy self-sufficient process to produce 10,000 t per year of biodiesel using supercritical methanol and propane as cosolvent. A suitable heat-integrated design of the process was performed using a part of the produced biodiesel (7.4 wt.%) to achieve the energy selfsufficiency for the plant. The simulations were performed by Aspen Plus considering a continuous multitubular reactor operating at 280 °C, 128 bar, methanol-to-oil molar ratio of 24, propane-to-methanol molar ratio of 0.05, and a residence time of 9.7 min. The use of flash separators instead of distillation columns for methanol and propane separation contribute significantly to reduce the energy requirement, with respect to other studies. Final products are high purity of methyl esters (99.8%) and glycerol (96.4%). The economic analysis of the biodiesel plant resulted in a break-even point for the biodiesel selling price of 0.479 €/kg (at 10% rate of return), so the process can compete with the existing alkali and acid catalyzed processes. Keywords Biodiesel, supercritical methanol, propane, simulation, techno-economic, energy integration 2 1. Introduction Biodiesel is a type of efficient renewable energy resource that can be an alternative to fossil energy. Sustainability of biodiesel production depends on a technology proficient to process feedstock with high content of free fatty acids and moisture, such as waste cooking oil. This way, non-catalytic biodiesel production by single-step transesterification with supercritical methanol is highly efficient, because only relatively short processing time and simple postprocessing for the product are needed [1]. The optimal conditions are 350 °C, 20 MPa and a methanol-to-oil molar ratio of 42:1, which lead to a large energy consumption in the reactants pre-heating and recycling steps is required, as a result of the high alcohol-to-oil molar ratio [2]. However, the high energy cost in the supercritical process could be compensated by catalyst costs in the alkali-catalyzed process and by less restrictive feedstock (water and fatty acids contents) . Current research has been focused on modifying the process, adding different heterogeneous catalysts or cosolvents, aimed at optimizing the operational parameters conditions (temperature, pressure and alcohol-to-oil molar ratio), thus leading to lower investment cost (materials) and energy cost (heat needed). Thus, the system can be modified by a two-step process, which involves subcritical hydrolysis and supercritical esterification [3], but the two-step process is more complicated than the single-step process, and hence its possible industrial implementation. Likewise, subcritical conditions have been considered to reduce not only pressure and temperature (150 ºC and 20 MPa) but also the methanol-to-oil molar ratio to 15:1, but such operating conditions require the use of effective heterogeneous catalyst [4], which leads to stability problems and an increase in the operating cost. On the other hand, the addition of a cosolvent, such as propane or CO2, reduces the alcohol-to-oil molar ratio to half. This is because the critical point of the reaction mixture depends on the alcohol-to-oil molar ratio, so the selected value for this ratio will reflect the operating temperature and pressure [5], 3 so the critical temperature decreases with the methanol-to-oil ratio and with the propane concentration. The addition of a small amount of gaseous cosolvent dramatically decreases the critical point of the reaction mixture allowing the use of milder operating parameters [6]. Thus, propane has been used as cosolvent at 280 ºC and 12.8 MPa for a molar ratio of methanol-to- oil as 24:1 and a molar ratio of propane-to-methanol of 0.05:1, while maintaining a high reaction rate [7]. There are no many papers focused on the industrial plant design and simulation of biodiesel production that includes a techno-economic evaluation using low-grade feedstocks, such as waste cooking oil. In such studies, biodiesel production under supercritical conditions was performed by different software as Aspen Plus [8-10], Aspen Hysys [11, 12] or SuperPro Designer [13]. Likewise, a few works on economic feasibility are found in the literature, in which biodiesel production under supercritical conditions resulted cost-competitive [11, 14, 15] or not [13], with respect to the existing alkali and acid catalyzed homogenous processes. Among all the studies found, some of them used thermodynamic methods, which are not suitable for all the process conditions [9, 11, 14], so the results may not be reliable. In addition, other studies [10] do not include the required purification of FAME and methanol recovery to meet the international standard of biodiesel, thus reducing apparently the energy consumption. In addition, the heat-integration is normally missing, as in [11-13, 15], or insufficient (any other, to our knowledge). However, this is a way to improve the process performance and economics. Furthermore, previous works have not tried to develop an energy self-sufficient process while performing the design and cost estimation of equipment, as well as obtaining a break-even point for the biodiesel selling price to be compared with commercial selling prices in order to show the feasibility of the process. This is the aim and scope of this paper, where the application of a more efficient process design is proposed while remaining the original simplicity and fulfilling the international standard of biodiesel (EN14214), which requires over 96.5% esters 4 content. This way, the purpose of achieving a lower energy consumption implies a better energy management and improves not only the economy of the process but also its environmental performance. 2. Methodology This section provides a conceptual design of an energy self-sufficient process to produce biodiesel under supercritical conditions, as well as the methods used to design, simulate and techno-economically assess a biodiesel plant with an annual production of 10,000 t during 330 days. 2.1. Process design and simulation Procedures required for the simulation of a process include the definition of the chemical components, selection of the thermodynamic methods and choice of operating conditions, as well as appropriate equipment and specifications. From an industrial point of view, a continuous process for biodiesel production with supercritical methanol is more appropriate than a batch process. Likewise, methanol is the most used alcohol in the world for biodiesel production because it is cheaper and more reactive than ethanol or other alcohol. Biodiesel production process by single-step supercritical transesterification was selected as the simplest non-catalytic supercritical transesterification process. Fig. 1 illustrates a flowsheet of the process performed in Aspen Plus v8.4. The selected feedstock was waste cooking oil. Calculations of the process were carried out by Aspen Plus using triolein (C57H104O6), as a model compound representative of triglycerides. This latter along with methanol and propane are the process feed. Likewise, the products are biodiesel, represented by methyl oleate (C19H36O2), and glycerin. The methanol is fed in excess to the reactor at a rate of 24 moles per mole of oil. Since the gaseous propane cosolvent can be easily separated from the products by expansion and be directly recycled, and thus almost 5 recovered, only a make-up propane of less than 1.0 kg per week would be necessary. This is very low for the steady-state process is very low. Thus, the process is fed continuously with fresh methanol and oil, while the propane is fed discontinuously, as a batch step, which was not to added to the flowsheet of the continuous process to avoid any complication in the simulation of it. The influence of it on the overall simulation and results is negligible. Fresh feed of raw materials is carried out under standard conditions (25 °C and 1 bar). Oil and methanol are pumped to 5 bar (pumps P1 and P2), and the oil feed is preheated to 40 °C by the HE1 exchanger. Both oil and methanol are mixed with the subsequent recirculation, which is also at 5 bar. Then, the mixture is pumped to 128 bar (P3) and preheated to 280 °C by a heat exchanger train (HE2, HE3, HE4, HE5). In the first three HE, the hot streams are process streams but, to achieve 280 ºC at the reactor inlet, it is necessary to use another energy source. Thus, it was chosen to burn a part of the produced biodiesel in a combustion chamber, and use the combustion gases (HE5) as a source of own energy in the production process in order to achieve an energy self-sufficient process. An optimal use of energy of the flue gas (stream CB1) involves to transfer it, first, in those points with the highest temperature and, lastly, in those locations with lower temperature. Thus, they provide first the required energy in the reboiler of the distillation column (T1), then in the reactor (R1), and after those preheat the reactor feed (HE5) and the product (HE6), once expanded in the valve and before entering the flash separator F1. The reaction takes place in a multitubular reactor (R1), where a power law kinetic model (Arrhenius) was adopted to model the plug flow reactor. Since the final products of the overall reaction are methyl esters and glycerin, it was used a simple mathematical model for this reaction, ignoring the intermediate reactions [16]. As kinetic parameters, an activation energy of 38,482 kJ/kmol and a reaction rate constant of 7ꞏ10-3 s-1 were used [14, 16]. The reaction rate is defined as irreversible (due to the high alcohol-to-oil molar ratio the reaction is very shifted 6 towards the product) and of first order in triglycerides (represented by triolein). The reaction rate increases considerably when propane is added to the supercritical system, and a similar value of that is used at the milder conditions with respect to those tested by [16], in which conversion was even somewhat lower. This way the kinetic model is based on overall reaction and it assumed that kinetic parameters provided by [16] may be used as an approach at milder conditions (for which there is not kinetic studies) used in this work. By a sensitivity analysis, it was obtained a reactor volume of 1.3 m3 to achieve maximum conversion (98.3%), requiring 9.7 min of residence time. The unreacted products can contain tri-, di- and mono-glycerides, and even some other compounds, formed as intermediate products, but they are probably present only in traces when almost completion of reaction, practically achieved at the reactor outlet. A multitubular reactor was selected as more suitable for the high pressure due to the higher mechanical strength of the tubes design. Once fixed the reactor volume, for the calculus of the diameter and length of the tubes, a Reynolds number of 10,000 was chosen to ensure a turbulent flow regime, thus avoiding a radial temperature variation and a dissimilar temperature profile along the reactor. Therefore, the reactor has three tubes with an internal diameter of 20 cm and a length of 13.8 m. Optimization of the methanol recovery is crucial to achieve a competitive overall process. One of the novelties implemented in this work with respect to other studies [8, 10] is the use of two adiabatic flash separators instead of distillation columns to recover methanol. Despite a distillation column is required to remove unreacted triolein from the final product, preseparation of most of the methanol in flash separators minimizes the flow-rate processed by the distillation column. This decision allows the reduction of investment and operating costs. Therefore, 99.0% of the excess methanol fed to the reactor, as well as the entire propane, are recovered using two flash separators (F1, F2). The first flash drum operates at 5 bar and 121 ºC. The preheating to 121 °C is performed in the HE6 by flue gas (entering at 323 °C). Methanol 7 and recovered propane are mixed with other reagents recirculated, while liquid stream is conducted to a second flash separator (F2), which operates at vacuum (0.02 bar). The feed to the separator F2 preheats the reactor feed in HE2, thus cooling down to 101 °C. The operating conditions for flash separators were determined to maximize the amount of recovered methanol, with restriction of separating a minimum of 99.9% methanol plus propane. The stream of methanol and propane recovered in F2 is compressed to 5 bar in a multistage compressor (C1) with intercooling. The stream leaving the compressor is then mixed with the vapor phase recovered in the flash separator F1, and the resulting stream preheats the reactor feed in HE3, then the feed to the distillation column (HE8), and finally the fresh oil feed (HE1). After this energy integration, stream leaving HE1 is at 107 °C and with 48.3% of vapor phase, which cannot be mixed and pumped with the rest of the feed. To make this possible, the heat exchanger HE10 is added, where a condensate is obtained at 41 ºC and 5 bar, by using cooling water. This pressure value is needed for condensing all the vapor phase. The liquid stream obtained at F2 is separated by the decanter S1 in a glycerin-rich phase and another rich in methyl esters. The glycerin-rich phase already complies with the relevant specifications needed for its sale on the market, while the phase rich in methyl esters needs to be purified to meet international standards. Unlike other studies [8-10], glycerin is obtained directly from the decanter with a high purity (96.4%), without needing a further purification for many applications, and it may be directly sold. However, the purity is slightly less than the 98% required for sale as technical grade glycerin. Purification of biodiesel consists mainly of the separation of the triolein from the product, and it is carried out in a distillation column (T1) operating under vacuum, with a reflux ratio of 2 and a partial condenser. The pressure in the reboiler and the condenser is 0.2 and 0.1 bar, respectively. The feed to the column is preheated by HE7, HE8 and HE9, in order to reduce the input of energy required in the reboiler. The stream leaving the column bottom contains more 8 than 99.9% of the triolein fed to the column. At the top, a vapor phase distillate stream containing only methanol and a liquid stream of biodiesel with 98.86% purity are obtained. This latter has a maximum methanol content of 0.12 wt.% and a glycerin content of 0.02 wt.% that meet the specifications UNE-EN14214. As above mentioned, the reactor model does not include intermediate compounds, but they are probably present only in traces at the reactor outlet, and likely they do not affect the final quality. The design was performed using first the module DSTWU, which applies the Winn-Underwood-Gilliland shortcut design calculations, and then the Radfrac method, which is a more rigorous distillation method to refine the previous results. As can be seen in Fig. 1, an energy integration was performed using a set of heat exchangers to preheat the reactor feed (HE2, HE3, HE4), column feed (HE7, HE8, HE9) and the feeding of fresh oil (HE1), following a strategy that minimizes the use of external utilities and maximizes internal heat flows into the system. Thus, small temperature-driving forces must be achieved, by using countercurrent flow and small temperature approaches at the ends of the exchangers. A ‘from inside to outside’ approach [17] was used to design the heat exchanger network, in such a way that the internal streams close to each other at high thermal levels are the first to contact each other (the high-temperature hot streams heat the high-temperature cold streams) and the external more separated streams at low thermal levels are the last to contact each other (the low-temperature hot streams warm up the low-temperature cold streams), by taking the reactor and the combustion chamber as reference central points. This way, the distribution of exchanger duties was assessed and maximum thermal energy could be recovered and used in the own process. The UNIQUAC model is not feasible to simulate equipment operating at high pressures since it is limited to low pressures, so it has been used only in zones of low pressure (up to 5 bar). For the high pressure sections of the process, a few thermodynamic models based on 15 The column was designed with the aid of the Aspen module 'Sizing and Rating', obtaining a diameter of 1.2 m. The spacing between plates obtained was 0.6 m, to allow the access of an operator. The cost of the column was calculated by Capcost, using AISI 316 as material. The estimation of the cost of the transesterification reactor was difficult because it is a very specific equipment, and there are still no plants that use this process on a commercial scale. Since it is a multitubular unit, the reactor was treated as a shell-and-tube exchanger for estimating its cost. The cost was scaled again using the rule of 'six tenths', taking the area of heat exchange (m2) as capacity factor [22], and updating by the CEPCI index. The waste oil should be filtered before entering the reactor, which is not included in the flowsheet of the process, but the cost of a filter press was estimated from equipment vendor [30], by applying the rule of 'six tenths' from 25 tonnes/day. Finally, the cost of combustion chamber was obtained from a costing tool [31] by applying the rule of 'six tenths' with the thermal power (kWt) as the capacity of the equipment to be purchased from 2000 kWt to 1000 kWt, and updating to the current date by the CEPCI index. Table 5 shows the estimated investment and operating costs, and Fig. 2 shows a breakdown of the total annual costs of biodiesel production, by annualizing the total capital costs at a 7% rate over the lifetime of the project (20 years). As can be seen, variable operating costs are the most significant mainly due to the waste cooking oil price (about 75% of the total annual cost). 3.2.2. Cash flow analysis On the other hand, according to the cash flow analysis made to determine the minimum selling price of a mass unit of biodiesel, this should be sold at 0.479 €/kg to fulfil all the requirements, indicating that the process becomes profitable with a biodiesel selling price of over this break-even point. Fig. 3 shows the variation of the cumulative non-discounted and discounted (present) cash flow over the lifetime of the plant. At the beginning, the cash flow is 16 negative due to the money paid (50% TIC). When the process is ready to start, there is an additional quantity of money to consider apart from the rest of capital costs (50% TIC), which is the working capital (5% TIC), so this is the point with the most unfavourable cash flow. However, when the production begins in the second year, money enters from sales (of biodiesel and glycerin) and the investment starts to be recovered at an internal rate return of 10%. After fifteen more years, the cash flow is recovered slower because the depreciation of capital goods is no longer accounted for. The payback period is the number of years to recover the initial investment of a project (discounting the working capital) from the cash flows that the project produces. It should be less than the project life to be profitable, and it is seen that the simple (non-discounted) payback period and discount payback period are 7.5 and 14.8 years, respectively, from the beginning of the production. The non-discounted cash flow accumulates, moving from negative to positive, and a positive final cash flow is observed beyond the payback period up to reach a value of 3.54 M€. The base-case analysis was performed assuming that no money is borrowed (100% equity financing). A more likely scenario is the debt financing of some fraction of the capital. Thus, by using fixed annuities, if 100% of the required capital was borrowed at 7% real interest with a 15-year repayment term (depreciation time), a comparison with the equity financing case indicates that the inclusion of debt financing would reduce the required selling price of biodiesel to 0.473 €/kg, and the uniform capital recovery factor would be 10.98 %. Despite the cost of financing the investment, the minimum selling price of biodiesel is approximately 1.25% lower. This is due to the effect called 'amplifying financial leverage', since the return on investment (10%) is higher than the interest rate on debt (7%). Therefore, in spite of the finance costs linked to debt financing, NPV increases at a given rate of return when access to credit is available at a not very high interest rate. 17 Estimates, based on the methodology used [19], have an expected accuracy of roughly ±30%. However, from Fig. 2, it is expected a moderate influence of the investment costs. Thus, if the TIC was one million of euros the selling price of biodiesel would decrease only in 2-3 cents of euro. Equally, to get a selling price of 0.50 €/kg, the TIC should be 50 % higher than that obtained for the base-case. Likewise, this also occurs with fixed operating costs or some variable operating costs such as methanol, process water or electricity, so the variation would be of ±1.5 cents of euro. Therefore, it is clear that the sensitivity analysis will show a significant effect of the waste oil cost on the biodiesel break-even price. Fig. 4 shows the influence of changes applied on waste cooking oil purchase price from -30% to +30% (with steps of 10%) around the base case (0%) on the biodiesel selling price, so this varies from 0.398 to 0.586 €/kg, which involves a variation of -8.1 and +10.7 cents of euro. It should be kept in mind that 7.4% biodiesel is internally used for energy self-sufficiency, and this fraction is relatively low thanks to the efficient heat integration performed in the process. 3.2.3. Techno-economic feasibility The suitable design of the process (optimal energy integration, milder operating conditions due to the use of a cosolvent, as well as the efficient and simple separation operations) along with the waste cooking oil cost, which is normally 2.5–3.5 times cheaper than that of vegetable oils [32, 33], support the feasibility of the process proposed in this paper. Thus, from the analysis performed and considering that, since biodiesel is usually obtained by conventional process (mostly alkaline catalytic process) with market prices varying currently to be greater than 0.75 €/kg [34-37], the biodiesel production under supercritical conditions seems to be a very promising technology with competitive production costs and a minimum selling price of 0,482 €/kg, thus allowing a considerable profit margin. Furthermore, the wholesale price of (fossil) diesel is between 0.390 €/kg (for resale) and 0.427 €/kg (to end users) [38], which are values close to that obtained in this study. In addition, as compared with conventional (fossil) 18 diesel, biodiesel production reduces significantly the amount of CO2 emitted to the atmosphere and, in contrast to other biodiesel production technologies that use vegetable oil, the use of waste oil as a raw material adds value to the process studied in the present paper because of the valorization of a waste, apart from facilitating waste management since this would have the biodiesel production plant under supercritical conditions as destination. Finally, it should be expected a reduction of the biodiesel selling price in a full-scale industrial plant, with a higher capacity (about 10 times more), by applying the economy of scale concept. 4. Conclusions A biodiesel production process using supercritical methanol and propane as a cosolvent has been designed and rigorously simulated by Aspen Plus, using the PSRK EOS as thermodynamic method, aimed at optimizing the power consumption by heat integration and achieving an energy self-sufficient process in terms of thermal energy (electricity consumption 79 kWe for an annual waste oil feed of 10,000 tonnes). This implies to burn 7.4% of produced biodiesel. The transesterification takes place at 280 °C and 128 bar for 9.7 min. Biodiesel and glycerin purity were 98.86% and 96.40%, respectively. The use of flash separators instead of distillation columns for methanol and propane separation contribute significantly to reduce the energy requirement, in relation to other studies. Besides, the energy scheme also involves the use of the cooling water for cogeneration (e.g., hot water for district heating). 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[34] S&P Global Platts. <http://www.platts.com.es/price-assessments/agriculture/european-biodiesel> (Accessed by Nov 13, 2016) [35] Piedmont biofuels, <http://www.biofuels.coop/price-of-biodiesel> (Accessed by Nov 13, 2016) [36] Biodiesel magazine, <http://www.biodieselmagazine.com/articles/586737/price-gap-between-diesel-biodiesel-widens-in- germany> (Accessed by Nov 13, 2016) [37] U.S. Department of Energy, Energy Efficiency & Renewable Energy, “Clean cities: Alternative fuel price report”, July 2016. <http://www.afdc.energy.gov/uploads/publication/alternative_fuel_price_report_july_2016.pdf> (Accessed by Nov 14, 2016) [38] U.S. Energy Information Administration <https://www.eia.gov/dnav/pet/PET_PRI_REFOTH_DCU_NUS_M.htm> (Accessed by Nov 14, 2016) Tables and Figures Captions Table 1. Mass and heat balance of the process streams. Table 2. Overall mass and heat balance in the operation units. Table 3. Comparative with other technologies/studies. Table 4. Purchase costs and main characteristics of operation units. Table 5. Estimated investment and operating cost, assuming 330 days per year. Fig. 1. Flowsheet of the process. Fig. 2. Breakdown of annual costs of biodiesel production. Fig. 3. Cumulative non-discounted and discounted cash flow diagrams over the lifetime at 100% equating funds. Fig. 4. Sensitivity analysis of the cost of biodiesel production versus waste oil purchase price. Table 1 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 Temperature (ºC) 25.0 25.0 41.0 25.3 40.0 25.2 35.7 41.8 55.6 129.7 278.0 280.0 280.0 138.6 110.2 121.0 121.0 101.0 29.6 33.8 Pressure (bar) 1.0 1.0 5.0 5.0 5.0 5.0 5.0 128.0 128.0 128.0 128.0 128.0 128.0 128.0 5.0 5.0 5.0 5.0 0.02 0.02 Mass flow-rate (kg/h) Methanol 0.0 146.2 967.1 0.0 0.0 146.2 1113.3 1113.3 1113.3 1113.3 1113.3 1113.3 977.2 977.2 977.2 977.2 275.5 275.5 10.0 4.8 Methyl esters 0.0 0.0 0.1 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1267.0 1267.0 1267.0 1267.0 1266.8 1266.8 1266.8 0.0 Triolein 1282.0 0.0 0.0 1282.0 1282.0 0.0 1282.0 1282.0 1282.0 1282.0 1282.0 1282.0 20.6 20.6 20.6 20.6 20.6 20.6 20.6 0.0 Propane 0.0 0.0 76.7 0.0 0.0 0.0 76.7 76.7 76.7 76.7 76.7 76.7 76.7 76.7 76.7 76.7 2.9 2.9 0.0 0.0 Glycerin 0.0 0.0 0.8 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 130.6 130.6 130.6 130.6 130.0 130.0 129.8 129.6 Overall mass flow-rate (kg/h) 1282.0 146.2 1044.8 1282.0 1282.0 146.2 2472.0 2472.0 2472.0 2472.0 2472.0 2472.0 2472.0 2472.0 2472.0 2472.0 1695.7 1695.7 1427.2 134.4 Vapor fraction 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.0 1.0 1.0 0.0 0.2 0.6 0.0 0.0 0.0 0.0 Liquid fraction 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 0.0 0.0 0.0 1.0 0.8 0.4 1.0 1.0 1.0 1.0 Enthalpy (kW) -50.3 -17.2 -116.6 -50.3 -49.7 -17.2 -183.4 -184.5 -183.1 -174.9 -154.4 -154.1 -151.7 -172.2 -170.2 -161.9 -91.1 -92.6 -66.1 -15.5 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 Temperature (ºC) 33.9 33.8 33.8 88.2 98.1 106.3 287.8 287.9 108.1 100.6 100.7 43.8 100.6 29.6 199.8 121.0 139.7 108.1 107.9 107.6 Pressure (bar) 1.0 0.02 0.2 0.2 0.2 0.2 0.2 1.0 1.0 0.1 1.0 1.0 0.1 0.02 5.0 5.0 5.0 5.0 5.0 5.0 Mass flow-rate (kg/h) Methanol 4.8 5.2 5.2 5.2 5.2 5.2 0.0 0.0 0.0 1.5 1.5 1.5 3.7 265.4 265.4 701.7 967.1 967.1 967.1 967.1 Methyl esters 0.0 1266.8 1266.8 1266.8 1266.8 1266.8 34.1 34.1 34.1 1232.7 1232.7 1232.7 0.0 0.0 0.0 0.1 0.1 0.1 0.1 0.1 Triolein 0.0 20.6 20.6 20.6 20.6 20.6 20.6 20.6 20.6 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Propane 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 2.9 2.9 73.8 76.7 76.7 76.7 76.7 Glycerin 129.6 0.2 0.2 0.2 0.2 0.2 0.0 0.0 0.0 0.2 0.2 0.2 0.0 0.2 0.2 0.6 0.8 0.8 0.8 0.8 Overall mass flow-rate (kg/h) 134.4 1292.8 1292.8 1292.8 1292.8 1292.8 54.7 54.7 54.7 1234.4 1234.4 1234.4 3.7 268.5 268.5 776.2 1044.8 1044.8 1044.8 1044.8 Vapor fraction 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.0 1.0 1.0 1.0 1.0 0.5 0.5 0.5 Liquid fraction 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 0.0 0.0 0.0 0.0 0.0 0.5 0.5 0.5 Enthalpy (kW) -15.5 -50.6 -50.6 -48.2 -47.7 -47.3 -1.6 -1.6 -2.0 -45.2 -45.2 -47.6 -0.4 -26.5 -25.4 -70.8 -96.2 -104.3 -104.8 -105.4