Abstract
This work addresses the process and economic performance of the production of gasoline and diesel range fuels from urban sewage sludge. The overall production route involves direct conversion of the sewage sludge to an intermediate oil phase, so-called biocrude, via hydrothermal liquefaction at near-critical water conditions and further upgrading of the biocrude based on conventional refinery processes. The overall mass and energy yields of combined naphtha and middle distillate from sewage sludge on dry basis are approximately 19 and 60%, where the naphtha fraction represents about 45% of the total, with a minimum fuel selling price ranging between 2.4 and 0.8 €/liter assuming full investment in both the biocrude production and upgrading plant with sewage sludge feed capacities in the range of 3 to 30 dry-ton/day. If existing equipment at refinery can be used for upgrading of the biocrude, the minimum fuel selling price can be reduced by approximately 7%.
1. Introduction
The use of advanced biofuels has been included as part of a wider global strategy for reducing GHG emissions in the transport sector.1 The role of advance biofuels is particularly relevant for achieving the decarbonization targets in the heavy-duty transport sectors, i.e., marine, aviation, and long-haul road transport, where the use of other alternative renewable energy carriers like electricity or hydrogen is not feasible. The total demand of liquid biofuels worldwide is estimated to be approximately 27% of the total fuel needed by 2050, with a similar share of about 25% to be reached within the European Union by 2030. However, despite the high short-term market demand, progress in commercialization of advanced liquid biofuels is still limited. One main reason is the high capital and operating costs of these technologies2 that require large scales to reach reasonable production costs. The need of larger plant capacities significantly increases the cost of feedstock transport3 as well as the risk of assuring a continuous supply of the feedstock throughout the lifetime of the plant. These reasons lead to financial risks when evaluating the viability of commercial projects, which hinder the commercial realization.
This work addresses the economic feasibility of producing liquid biofuels from urban sewage sludge. The overall conversion route considered the decentralized conversion of sewage sludge to biocrude via hydrothermal liquefaction (HTL) at near-critical-water conditions and further upgrading of the biocrude to liquid biofuels based on conventional refinery processes. Therefore, the analysis explores two important strategies for improving the overall economy of producing advance liquid biofuels, i.e., to lower the feedstock supply cost by utilizing low-grade organic waste and to reduce capital cost by utilizing of existing petrochemical infrastructures for refining of organic intermediates to marketable biofuels. Sewage sludge has been considered in this analysis as a model feedstock representing renewable urban waste. Sewage sludge is an abundant feedstock, with an annual production in the European Union being approximately 10 million tons4 in 2019. At present, the commercial disposal of sewage sludge includes the direct spreading in soils,5,6 anaerobic digestion for production of biogas,7 thermal conversion for production of heat and power,8−10 and co-combustion in kilns for cement production.11 Due to its high moisture content, typically above 80% wt., the transport and direct combustion of sewage sludge is energetically unfavorable. It also causes severe fouling problems in boilers due to alkali and phosphorous contents.12 Anaerobic digestion of sewage sludge achieves high methane yields in the range of 230–430 Nm3 CH4/ton volatile solid13 depending on the pretreatment methods and digestion conditions. However, it requires typically large conversion times leading to high reactor volumes and thus high capital investment.10 In this context, hydrothermal liquefaction (HTL) technology exhibits several advantages for the conversion of wet organic waste fractions to energy.14−16 Under near critical or near-critical conditions, liquid water exhibits high ionic product.17,18 These conditions enhance the decomposition of macromolecular constituents of the feedstock by ionic reactions taking place in the liquid water with formation of a high-energy density oil phase,17 or so-called biocrude. Depending on the operational conditions and feedstock characteristics, conversion times in the HTL process are typically below 30 min,15 which is about three orders of magnitude lower than the conversion to biogas in anaerobic digestion. Moreover, the chemical and physical characteristics of the biocrude produced from HTL make it compatible with fossil-based crude oil for upgrading to liquid fuels in conventional refinery processes.19
The techno-economic analysis of the production of liquid biofuels from sewage sludge based on HTL technology has been addressed in the literature.19,20 These two studies by the same author considered decentralized production of biocrude with centralized upgrading at refinery. The former study assumed the same design and product yields for the HTL process as considered earlier21 for the conversion of algae showing a minimum biocrude selling price (MBSP) of $2015 31.6/GJ and a minimum fuel selling price (MFSP) of $2015 40.75/GJ for a base conversion capacity of 100 dry ton/day covered by 10 identical HTL plants for production of biocrude. The later study by the same author performs a more detailed evaluation of the HTL and upgrading processes for several wet waste fractions based on measurements of the product yields and composition from engineering-scale and bench-scale experiments. The results from this work on the conversion of sewage sludge generated at waste-water treatment (WWT) plants showed values of the MBSP per unit volume and energy of 0.67 $2021/liter and 19.5 $2021/GJ, respectively, and the MFSP per unit volume diesel equivalent and per unit total hydrocarbons products energy of 1.02 $2021/liter and 29.8 $2021/GJ, respectively. This analysis assumed a decentralized HTL plant with a constant capacity of 110 dry ton/day and centralized upgrading at refinery also with a constant capacity of 30.4 m3/day. The process design of the HTL plant considered treatment of the HTL process water by aqueous catalytic upgrading followed by ammonia stripping before disposal back to the WWT plant. Other investigation addressing the techno-economics of hydrothermal liquefaction for the conversion of wet organic waste of biocrude22 has shown MBSP values in the range of 22 and 41 $2020/GJ. The analysis presented in this paper contributes to the topic of techno-economics of HTL biofuel production from sewage sludge in several relevant aspects. It provides a detailed description of the process design that includes the main process and auxiliary equipment, allowing a realistic estimation of the capital cost. The analysis has been performed through parametric models for evaluating the main material and energy flows from experimental results as well as for evaluating the main size of equipment. This parametric approach has been used for scaling up the process and evaluating the economy as a function of the production capacity.
2. Description of the Process Design
2.1. Production of Biocrude from Sewage Sludge
The process design for the overall conversion of sewage sludge to biocrude is shown in Figure 1. The raw feedstock is stored in a buffer tank and transported by a screw conveyor into the slurry preparation tank where it is mixed with the catalyst (K2CO3), the base (NaOH), and a fraction of the concentrate from the HTL water treatment. The slurry preparation tank is designed as a stirred tank to ensure uniform mixing of the slurry with a heating jacket to increase the temperature of the slurry for reducing the viscosity and thus improving the pumpability. The slurry is discharged from the stirred tank by a pump that feeds the main pressurization pump, which increases the pressure of the slurry to the HTL conditions. The main pressurization pump is designed as a single piston low stroke reciprocating pump. The pressurized slurry is heated to the HTL conditions in a U-tube heat exchanger and then fed into the HTL reactor. The process design considers one or several parallel HTL reactors depending on the capacity of the biocrude production plant. The HTL reactor is equipped with a heating jacket and dimensioned to achieve the required residence time to complete the liquefaction of the slurry. The HTL product is a multiphase flow consisting of a binary liquid mixture of aqueous and oil phases with dispersed non-dissolved solids and a gas phase. The product from the HTL reactor is directly cooled down in a heat exchanger before depressurization. Heating of the slurry and cooling of the HTL product are performed in a closed loop using a thermal fluid that recovers the heat from cooling of the HTL product. The cooled liquefaction product is partially depressurized in capillary columns to an intermediate pressure of about 30 bar-g and then taken to a gas separator where the gas phase is separated from the oil, aqueous, and solid phases. These three phases leave the separator as an oil–sand emulsion and stored in a buffer tank for further separation in a two-stage centrifugation system. The first centrifuge separates a fraction of the water. The remaining oil sand emulsion from the first centrifuge is mixed with acid, methyl ethyl ketone (MEK), and a condensed light oil stream for breaking the oil/water chemical binding and then fed into the second centrifuge where the oil, solid, and aqueous phases are separated. MEK is here used only during startups of the plant since the condensed light oil effectively can be effectively used as a solvent to break up the oil–sand emulsion. The solids from the second centrifuge are collected in a hopper and transported by a screw conveyor to a silo. The oil from the centrifuge, here also called biocrude, is stored in a tank for further delivery. The aqueous phase from the second centrifuge is discharged into a flash tank. Compressed air is injected into the flash tank for stripping of the ammonia, light oils, and the MEK solvent. The gas stream from the flash tank is further cooled for condensation of the light oils and MEK, which are recirculated and mixed with the oil–sand emulsion before the second centrifuge, and the remaining gas stream is used as a combustion air in the HTL gas treatment. The HTL process water from the flash tank is stored in a buffer tank before treatment.
Figure 1.
Schematic representation of the biocrude production process from sewage sludge by hydrothermal liquefaction.
Treatment of the process water is achieved by partial evaporation of the process water in a mechanical vapor recompression (MVR) unit. The MVR technology has been selected in this analysis since it is commercially available and has been successfully used by Steeper Energy at a demo scale for treating HTL water derived from woody biomass and at a pilot scale for treating HTL water derived from sewage sludge. The MVR system effectively utilizes the latent heat from the vapor for partial evaporation of the inlet process water. The concentrate from the MVR unit is partially recirculated back to the slurry preparation tank before the HTL process, and the remaining concentrate bleed is transported by a screw conveyor to the solid residue silo. The cleaned condensate from the MVR unit is further cooled and stored in a buffer tank before disposal. The gas separated from the HTL product is directly combusted in a burner to produce the heat required by the process. The calorific value and flammability of the HTL gas product are typically low due to the high CO2 concentration. Therefore, the HTL gas is co-combusted with natural gas to provide additional thermal power to cover the total heat demand by the plant and compensate the variability in the energy content in the HTL gas. The flue gas from the combustion process contains SO2 and NOx and requires further cleaning to fulfill the air emission regulations. The first step in the flue gas cleaning is a selective catalytic reduction (SCR) of the NOx where urea is used as a reducing agent. After NOx reduction, flue gas is further cooled down and taken into a dry scrubber for removal of acid gases SO2 with hydrated lime. The scrubber is equipped with bag filters to remove the reacted hydrated lime, which is removed from the external surface of the bag filters by pulsing compressed air, collected at the bottom of the scrubber, and transported by a screw conveyor into a silo. The heat from the flue gas is recovered to the thermal oil by two heat exchangers before the SCR unit and the scrubber. The main process design parameters for the biocrude production process are listed in Table 1.
Table 1. Process Parameters for the Biocrude Production Process Considered in the Analysis.
| process parameter | value |
|---|---|
| storage capacity sewage sludge (h) | 12 |
| slurry preparation temperature (deg. C)a | 120 |
| slurry preparation pressure (bar-a) | 5 |
| agitation at slurry preparation tank (rpm) | 150 |
| slurry dry matter concentration (% wt.) | 25 |
| slurry pH | 9 |
| base consumption in HTL (% wt. input wet feedstock) | 0.34 (NaOH) |
| catalyst consumption in HTL (% wt. input wet feedstock) | 0.14 (K2CO3) |
| HTL temperature (deg. C) | 350 |
| HTL pressure (bar-g) | 320 |
| HTL residence time (s) | 1500 |
| HTL heating rate (deg. C/s) | 2.0 |
| gravimetric separation temperature (deg. C) | 150 |
| gravimetric separation pressure (bar-g) | 30 |
| temperature at aqueous effluent flash tank (deg. C) | 100 |
| pressure at aqueous effluent flash tank (bar-a) | 1.2 |
| temperature at first centrifugation | 150 |
| aqueous/oil/ash content after first centrifugation | 0.5/2/1 |
| centrifuge rotational speed (rpm) | 9510 |
| citric acid consumption oil separation (mol/liter emulsion) | 0.013 [1] |
| MEK consumption oil separation (kg/kg oil)a | 1.0 [1] |
| MEK recovery (%)a | 98 |
| MVR operating temperature (deg. C) | 110 |
| MVR operating pressure (bar-a) | 1.013 |
| MVR concentration factor (% wt. water reduction) | 80 |
| recirculation of MVR concentrate to HTL process (%) | 80 |
| boiler pressure (bar-g) | –0.05 |
| excess air in combustion | 1.1 |
| combustion air supply pressure (kPa-g) | 1.3 |
| natural gas supply pressure (kPa-g) | 1.3 |
| SCR temperature (deg. C) | 390 |
| ammonia consumption SCR (kg/Nm3) | 0.43 |
| SCR, electricity consumption (kWh/Nm3) | 1.4 × 10–3 |
| SCR, catalyst lifetime (hours) | 40,000 |
| dry scrubber temperature (deg. C) | 140 |
| lime consumption at dry scrubber (g/Nm3) | 2 |
| thermal fluid | thermal oil |
| thermal fluid supply temperature (deg. C) | 400 |
| thermal fluid supply pressure (bar-g) | 15 |
| dry matter content in the leaching tank | 10% |
| operating temperature at HTL solids leaching tank (deg. C) | 95 |
| operating pressure at HTL solids leaching tank (bar-a) | 1.013 |
| acid concentration at HTL solids leaching tank | 0.4 M (H2SO4) |
During plant start.
2.2. Biocrude Upgrading to Naphtha and Middle Distillate
The process design of the biocrude upgrading is graphically represented in Figure 2. The raw biocrude is stored in a buffer tank before treatment. From the buffer tank, the biocrude is pumped, mixed with hydrogen, and heated before entering the guard reactor, where inorganic heteroatoms in the biocrude are reduced and adsorbed by the catalyst. Heating of the feed is performed in a heat exchanger with recovery of thermal energy from the hydrotreating product. The product from the guard reactor is mixed with hydrogen, heated, and taken into the hydrotreating reactor where the N, S, and O heteroatoms are reduced and separated from the oil. Heating of the feed before hydrotreating is performed in a fired heater using light hydrocarbons separated during fractionation of the hydrotreated oil as a gas fuel. The product from the hydrotreating reactor is cooled by and taken to a high-pressure (HP) three-phase gravimetric separator where a light sour gas and water are separated from the organic liquid phase. The organic liquid from the HP separator is heated and enters a second low-pressure (LP) stripper where light hydrocarbons and water vapor are separated and condensed. The process water streams from the HP and LP separation are disposed to the water treatment system at the refinery. The oil phase from the LP stripper is fed directly into a first distillation column for separation of naphtha, with boiling point in the range of 90–210 deg. C. The bottom distillate from the naphtha column is re-boiled and fed into a second distillation column for separation of middle distillate, with boiling point in the range of 210 and 310 deg. C. The heavy distillate residue from the diesel column is pumped, mixed with hydrogen, heated, and fed into a hydrocracking reactor where the heavy distillate is converted to naphtha and middle distillate, and the remaining S, N, and O contained in the feed are reduced and separated from the liquid organic product. The product from the hydrocracking is cooled and mixed with the hydrotreating product before phase separation. The light sour gas from the HP separation, containing non-reacted H2, CO2, NH3, H2S, and light hydrocarbons, is further treated in an absorption column where rich amine is injected for solubilization of CO2, NH3, and H2S. The H2 enriched gas stream from the amine column is mixed with make-up hydrogen, compressed, and recirculated back to the guard, hydrotreating, and hydrocracking reactors. The lean amine with dissolved gases is heated and pumped into a stripper where the gases are desorbed from the amine. The rich amine after the stripper is recirculated back to the absorption column prior cooling. The off gas from the stripper is disposed to the refinery for further treatment. Table 2 lists the main process design parameters used for the upgrading of the HTL biocrude to naphtha and middle distillate.
Figure 2.
Schematic representation of the biocrude upgrading process.
Table 2. Process Parameters for the Overall Biocrude Upgrading Process Considered in the Analysis.
| process parameter | value |
|---|---|
| storage capacity biocrude (h) | 12.00 |
| preheated biocrude temperature before pumping (deg. C) | 40.00 |
| guard reactor temperature | 290.00 |
| guard reactor pressure | 100.00 |
| guard reactor hydrogen consumption (Nm3/m3 feed) | 574 |
| guard reactor hydrogen reacted (% wt. feed) | 0.52 |
| guard reactor catalyst WHSV (kg/kg/h) | 0.40 |
| guard reactor catalyst lifetime (h) | 16000.00 |
| hydrotreating temperature | 360.00 |
| hydrotreating pressure | 100.00 |
| hydrotreating hydrogen consumption (Nm3/m3 feed) | 1747 |
| hydrotreating hydrogen reacted (% wt. feed) | 1.57 |
| hydrotreating catalyst WHSV (kg/kg/h) | 0.40 |
| hydrotreating catalyst lifetime (h) | 16000.00 |
| hydrocracking temperature (deg. C) | 370.00 |
| hydrocracking pressure (bar-g) | 100.00 |
| hydrocracking hydrogen consumption (Nm3/m3) | 2400.00 |
| hydrocracking hydrogen reacted (% wt. feed) | 5.4 |
| hydrocracking catalyst WHSV (kg/kg/h) | 0.50 |
| hydrocracking catalyst lifetime (h) | 16000.00 |
| temperature at three-phase high-pressure separator (deg. C) | 40.00 |
| pressure at three-phase high-pressure separator | 100.00 |
| temperature at three-phase low-pressure separator (deg. C) | 240.00 |
| pressure at three-phase low-pressure separator | 2.00 |
| inlet temperature at distillation column (deg. C) | 400.00 |
| pressure distillation column (bar-g) | 1.013 |
| naphtha cut-off temperature at distillation (deg. C) | 95 |
| diesel cut-off temperature at distillation (deg. C) | 210 |
| heavy-fraction cut-off boiling point (deg. C) | 310.00 |
| diesel range cut-off boiling point (deg. C) | 210.00 |
| naphtha cut-off boiling point (deg. C) | 80.00 |
| amine absorber temperature (deg. C) | 40.00 |
| amine absorber pressure (bar-g) | 30.00 |
| make-up amine consumption (% wt. feed gas) | 7.00 |
| heating amine plant (MJ/kg feed gas) | 0.272 |
| net cooling amine plant (MJ/kg feed gas) | 0.40 |
| electric power consumption amine plant (kWh/kg feed gas) | 46.50 |
3. Process Models and Analysis
3.1. Flow Properties
The process analysis considers a general multi-phase slurry flow structure, with physical properties evaluated as a function of the composition and properties of the phases. For flows with solid particle dispersion, it has been assumed that the particles are suspended and distributed uniformly in the liquid phase. Then, the effects of particle interaction on the molecular transport properties can be neglected. The apparent density and the specific heat and enthalpy of the flow are calculated, respectively, from ρ = ∑ ϕkρk, cp = ∑ ykcp, k, and h = ∑ Ykhk, where ϕk, yk, ρk, cp, k, and hk denote the void fraction, mass fraction, density, specific heat, and specific enthalpy of each phase. The effective viscosity of the slurry23 is evaluated from μ̅ = μlμr, where μl represents the viscosity of the liquid phase and μr denotes a relative viscosity dependent on the particle geometry and concentration. The relative viscosity of the slurry feed prepared from sewage sludge is calculated as a function of the total solids volume fraction using the correlation μr = 1 + ϕsA/(1 – ϕs/B), with the constants A = 3000 and B = 0.27 estimated from measurements of the total dynamic viscosity of sewage sludge slurries at 320 bar-g. The thermal conductivity is calculated as a function of the thermal conductivity of the solid ks and liquid kl phases from k̅ = klkr, where kr is the relative thermal conductivity of the slurry evaluated from kr = 1 + 3ϕsβ + 3ϕs2β2[1 + (β/4)(β + 29/4)/(5 + β)], with β = [(ks/kl) – 1]/[(ks/kl) + 2] denoting the solid to liquid thermal conductivity ratio. The physical properties of water have been evaluated using the IAPWS formulation.24 When the molecular composition is defined, evaluation of the enthalpy and specific heat has been calculated using the NIST database of thermodynamic properties of fluid systems.25
3.2. Hydrothermal Liquefaction
The conversion of the slurry in the overall hydrothermal liquefaction system, including phase separation, has been described in terms of the mass flow rate of the slurry entering the HTL process ṀslHTL, the yields of mass mkand chemical energy ekHTL, the atomic composition yk, iof the products after the phase separation, and the overall heat demand QthHTL. Here, the subscripts i and k denote the atomic composition and the oil (O), gas (G), solid (S), and aqueous (A) phases. The mass flow rate of the slurry entering the HTL process is calculated from Ṁsl= ṀFHTL(1 + yb+ ycatHTL), where ycatand ybHTLdenote the mass fraction of the catalyst and base required by the HTL process. The product yields and composition have been evaluated semi-empirically throughout the so-called transfer coefficients fi, kand fA, kHTLdefined as the distribution of the atomic mass and aqueous phase from the input slurry among the different products after phase separation. From these definitions, the mass yield and the dry atomic composition of the oil, solid, aqueous, and gas streams after phase separation are calculated, respectively, from
 |
1 |
yi, kHTL= mi, k/∑imi, k with mi, k = yF, DMyF, iHTLfk, i+ fA, kHTL(yF, DMfi, AHTL+ ycatyi, cat + ybHTLyi, b) for k = O, S, G and mi, A= (1 – fA, OHTL– fA, S)[yF, DMHTL∑iyF, ifi, AHTL+ (1 – yF, DM) + ybHTL+ ycat]. The energy yield has been evaluated from ekHTL = (HHVk∑iyk, iHTL)/(yF, DMHHVFHTL), where HHVk are the high heating value of each product phase. In this formulation, ṀFHTL, yF, DM, yF, iHTL, and HHVF denote the mass flow rate, the dry matter content, the dry atomic composition, and the dry high heating value of the sewage sludge. It has been assumed that both the base and the catalyst are fully dissolved in the water and remain unreacted during conversion in the liquefaction process. Also, the process water in the oil, solid, and aqueous phases from phase separation after liquefaction has been assumed to have the same chemical composition. Values of the coefficients fi, kHTL, fA, k, and ekHTLare shown in Table 4, evaluated using the measurements of the composition and calorific value of the HTL product phases shown in Table 3. In this table, the measurements of concentrations in the aqueous phase have been already reported by Sayegh et al.26 It has been assumed that the nitrogen and sulfur content in the gas phase is in the form of NH3 and H2S and that the composition of hydrocarbons in the gas phase is a mixture of methane, ethene, ethane, propane, methanol, ethanol, and acetone with molar distributions of, respectively, 0.42, 0.26, 0.01, 0.16, 0.05, and 0.06 based on the data reported by Jensen et al.17 on near-critical liquefaction of lignocellulosic biomass. The net rate of heat demand in the overall HTL system is calculated from Qth= ṀFHTL[(1 – yF, DM)(hwPS– hw) + yF, DMHTLcF, DM(THTL – T0) – ∑kmG(1 – yG, H2OPS)cp, k(THTL – Tk)], where (hwPS– hw) is the relative enthalpy of water between phase separation and ambient conditions, THTL is the operating temperature of the HTL process, and TkPSand cp, k are the temperature and the specific heat capacity on dry basis of each product from phase separation. It has been assumed that specific heat capacity of the dry matter in the sewage sludge and the solid and aqueous phases is the same and equal to 1.2 kJ/kg K. The specific heat capacity of the dry gas stream has been calculated based on the molecular composition shown in Table 3.
Table 4. Estimated Transfer Coefficients for the for Hydrothermal Liquefaction of Sewage Sludge and Woody Biomass Performed at Near-Critical Water Conditions at the Aalborg Pilot Plant.
| oil phase | aqueous phase | solid phase | gas phase | |
|---|---|---|---|---|
| water | 0.16% | 77.7% | 20.2% | 1.9% |
| chemical enthalpy | 73.38% | 17.14% | 4.57% | 4.87% |
| total dry matter | 28.4% | 23.1% | 38.8% | 28.4% |
| carbon, C | 60.0% | 28.5% | 5.0% | 6.5% |
| hydrogen, H | 44.0% | 40.0% | 12.0% | 4.0% |
| oxygen, O | 7.0% | 25.0% | 52.0% | 16.0% |
| nitrogen, N | 21.0% | 62.0% | 6.0% | 11.0% |
| sulfur, S | 24.0% | 14.0% | 7.0% | 55.0% |
| phosphorous, P | 0.5% | 7.5% | 92.0% | 3.1% |
| calcium, Ca | 0.0% | 6.0% | 94.0% | 0.0% |
| aluminum, Al | 2.5% | 11.0% | 86.5% | 0.0% |
| iron, Fe | 1.4% | 4.0% | 94.6% | 0.0% |
| magnesium, Mg | 0.1% | 2.0% | 97.9% | 0.0% |
| potassium, K | 2.0% | 75.1% | 22.9% | 0.0% |
Table 3. Experimental Measurements of the Oil, Aqueous, and Solid Phases Produced from Hydrothermal Liquefaction of Sewage Sludge at 320 bar-g and 350 deg. C at the Aalborg Pilot Planta.
| feedstock | oil phase | aqueous phase | solid phase | gas phase | ||||
| HHV | MJ/kg daf | 35.2a | MJ/kg db. | 14.4 | MJ/kg db. | 2.38b | MJ/kg db. | 7.1b |
| water | wt. % | 1.8a | wt. % | 91.4b | wt. % | 35c | wt. % | 60c |
| carbon (total organic carbon) | wt. % daf | 77.4a | g/L | 45.3 (35.4)a | wt. % | 17.3b | wt. % db. | 31.6b |
| hydrogen H | wt. % daf | 9.8a | g/L | wt. % | 2.1b | wt. % db. | 1.4b | |
| oxygen O | wt. % daf | 8.2a | g/L | wt. % | 3.7b | wt. % db. | 61.2b | |
| nitrogen N (N as NH4+) | wt. % daf | 3.7a | g/L | 11.6 (7.98)a | wt. % | 1.2b | wt. % db. | 2.45b |
| sulfur S | wt. % daf | 0.8a | g/L | 0.648a | wt. % | 0.48b | wt. % db. | 3.3b |
| phosphorous P | g/kg dry | 0.03a | g/L | 0.7a | g/kg | 78.1b | g/kg dry | 0.0 |
| calcium Ca | g/kg dry | mg/L | 31.37a | g/kg | 34b | g/kg dry | 0.0 | |
| aluminum Al | g/kg dry | 0.0739a | mg/L | 65.23a | g/kg | 13.8b | g/kg dry | 0.0 |
| iron Fe | g/kg dry | 0.11a | mg/L | 66.27a | g/kg | 64.2b | g/kg dry | 0.0 |
| magnesium Mg | g/kg dry | 1.84 × 10–3a | mg/L | 8.64a | g/kg | 10.6b | g/kg dry | 0.0 |
| potassium K | g/kg dry | 7.03 × 10–2a | mg/L | 10,054a | g/kg | 22.5b | g/kg dry | 0.0 |
| carbon dioxide CO2 | % vol. dry | 75.6a | ||||||
| carbon monoxide CO | % vol. dry | 0.7a | ||||||
| hydrogen H2 | % vol. dry | 4.2a | ||||||
| ammonia NH3 | % vol. dry | 8.5b | ||||||
| hydrogen sulfide H2S | % vol. dry | 0.4b | ||||||
| total hydrocarbons | % vol. dry | 10.6b | ||||||
Notes: aMeasured; bCalculated; cAssumed; The composition of hydrocarbons in the gas phase is assumed to be a mixture of methane, ethene, ethane, propane, methanol, ethanol, and acetone with molar distributions of, respectively, 0.42, 0.26, 0.01, 0.16, 0.05, and 0.06 based on data reported by Jensen et al.17 on near-critical liquefaction of lignocellulosic biomass.
3.3. HTL Water Treatment by Mechanical Vapor Recompression (MVR)
The HTL water treatment in the mechanical vapor recompression (MVR) unit is defined in terms of the concentration factor ηMVR, defined as the ratio between the input mass flow rate of process water and the mass flow rate of concentrate, the net heat demand Q̇thMVRand the net electric power consumption Ẇel. The mass flow rates of concentrate and cleaned water are then calculated from ṀcMVR= ṀA/ηMVR and ṀwWT= ṀA(1 – ηMVR)/ηMVR. The net heat demand is calculated from Q̇thMVR= ṀFmAHTL[yA, H2O(hwMVR– hw) + ∑k(1 – yA, H2OPS)cp, A(TMVR – TA)], where (hwMVR– hw) is the relative enthalpy of water between phase separation and the MVR unit, and THTL is the operating temperature of the MVR unit. The electric power consumptions are calculated from ẆelMVR= V̇AwelMVR, where welis the specific electric load per unit feed volume, assumed to be constant and equal to27 39.8 kWh/m3.
3.4. Combustion of the HTL Gas
The analysis of the overall process in the boiler includes the mass and energy flows and composition of the flue gas at the boiler outlet, the mass and energy flows of combustion air, and the net thermal power transfer to the thermal fluid. The main input process parameters in the boiler are the mass flow rates of HTL gas and natural gas, denoted by ṀGHTLand ṀNG, the excess air ratio for the combustion process λgCOMBdefined as the ratio between the total inlet combustion air and the stoichiometric air, the inlet temperature for the combustion air Ta, and the flue gas temperature at the boiler outlet TgCOMB. Natural gas consumption in the boiler is calculated from ṀNGLHVNG= [(QthHTL+ Q̇th) – ṀGHTLLHVG]/ϵth, where ϵth is the thermal efficiency of the boiler defined as the output heat transferred to the thermal fluid relative to the input energy to the boiler from the HTL gas and the natural gas, which is assumed to be constant and equal to 0.9. The overall combustion process is calculated from steady-state mass and energy conservation equations and assuming the overall chemical reaction CHaNbOcSd + νO2(O2 + 3.76N2) → CO2 + (a/2)H2O + νNONO + dSO2 + (3.76νO2 + b/2 – νNO/2)N2. Here, CHaNbOcSd represents the chemical formula for each of the combustible species j in the input fuel in the HTL gas and the natural gas, with a, b, c, and d representing the atomic molar composition of H, N, O, and S relative to C, and νO2 = 1 + c/2 – a/4 – d is the moles of air required for stoichiometric combustion of one mol of each combustible species. Then, the mass and energy flows of the combustion air at the boiler inlet can be written, respectively, as {Ṁair,COMB,Ḣair} = ∑jṀjCOMB(MWair/MWj)νO2, j(λg/xO2air){1, hair}, where j denotes each combustible species, νO2, j is the stoichiometric moles of oxygen required for complete combustion of one mol of j, MWair is the molecular weight of the air, and hairCOMBis the molar enthalpy of air evaluated at Tair. The mass flow rate of the flue gas at the boiler outlet is calculated by applying a mass conservation equation to the entire boiler from ṀgCOMB= ∑jṀjmg, jCOMB, where mg, jis the specific mass of flue gas produced from each combustible species, which is obtained from mg, jCOMB= 1 + (MWair/MWj)νO2, j(λg/xO2air). The O2, N2, H2O, CO2, and SO2 compositions in the flue gas are calculated using the conservation equations for C, N, H, O, and S from
 |
2 |
In this equation, it is assumed that all sulfur present in the flue gas is in the form of sulfur dioxide. The concentration of NO has been estimated to be 290 ppm based on experimental results28 from combustion of methane and CH4/NH3 mixtures in gas turbines, which shows a constant asymptotic value for NO concentration in the flue gas when the NH3 concentration in the gas fuel is above 5% vol. The energy flow rate of the flue gas at the boiler outlet is evaluated based on the mass flow rate and composition from ḢgCOMB= ∑jṀjmg, jCOMBhg, j, where the specific enthalpy hg, jCOMBis estimated from hg, j= ∑jyj, gCOMB(h̅j/Wj), where h̅j is the molar thermal enthalpy for each species evaluated at Tg.
3.5. Flue Gas Cleaning after HTL Gas Combustion
Cleaning of the flue gas from the combustion process involves reduction of NOx and removal of SO2. The reduction of NOx is performed in a selective catalytic reduction (SCR) unit, the overall process performance being defined in terms of the NO removal efficiency ηSCR, the total volumetric flow rate of ammonia consumed, V̇NH3SCR= V̇gxg, NOxCOMBλNH3, the consumption of catalyst V̇catSCR, and the overall electric power consumption Ẇel= V̇gSCRwel. Here, the parameter λNH3SCRrepresents the moles of ammonia consumed by the SCR per unit mol of NOx in the flue gas. The NOx reduction efficiency is assumed to be 92% at an operating temperature of 390 deg. C.29 The consumption of catalyst is evaluated considering a reference catalyst lifetime 32,000 h typically achieved in SCR units after natural gas boilers.29 The composition of the flue gas after the SCR is calculated assuming an overall NO reduction chemistry defined by the overall reaction NO + NH3 + (1/4)O2 → N2 + (3/2)H2O. Then, the flow rates of the different gas species leaving the SCR can be written as V̇G, i= V̇G, iCOMBfor i ≠ H2O, NO, NH3, N2, and O2 and
 |
3 |
The total flue gas mass flow and its mass composition can then be written as xjSCR= V̇j/V̇gSCR, with V̇g= ∑jV̇jSCR. The reaction heat for the NO reduction is assumed to be small compared to the to the total thermal enthalpy of flue gas, and therefore, the gas temperature variation in the SCR have been neglected. The removal of SO2 in the dry scrubber (DS) is evaluated in terms of the consumption of lime Ṁlime= λlimeDSV̇gxSO2, gSCR, the electric power Ẇel= V̇gSCRweland the SO2 removal efficiency ηSO2DSThe consumption of quicklime has been assumed to be proportional to the total inlet SO2 molar flow rates where λlimeis a fixed coefficient representing the unit mass of lime consumed per mol of SO2. The operating temperature in the dry scrubber is assumed to be 140 deg. C, above the dew point SO2. The overall chemistry in the dry scrubber is defined by the overall reaction SO2 + Ca(OH)2 → CaSO3 + H2O. It has been assumed that the CaSO3 and unreacted lime are separated from the flue gas stream in the bag filters. The mass flow of the different gas species leaving the scrubber are calculated from Ṁg, iDS= Ṁg, ifor i ≠ H2O, SO2, Ṁg, H2ODS= Ṁg, H2O+ ṀlimeDSyH2Oand Ṁg, SO2DS= (1 – ηSO2)ṀSO2SCR. The total flue gas mass flow and its mass composition can then be written as Ṁg= ∑jṀg, jDSand yg, j= Ṁg, jDS/Ṁg. Since the overall heat of the reactions between hydrated lime and acid gases is small in comparison to the heat absorbed by the evaporation of the excess water, then it has been neglected.
3.5.1. Guard and Hydrotreating Reactor Processes
The overall processes in the guard and hydrotreating reactors are defined in terms of the mass yields mkSand composition yi, kof the oil, gas, and aqueous phases at the reactor outlet and the required input hydrogen ṀH2S. Here, the superscript S denote the guard (GR) or hydrotreating (HT) processes, the subscripts k denote the oil (O), gas (G), and aqueous (A) phases, and the subscript i denotes the atomic composition. The input hydrogen to each process S is calculated from ṀH2= ṀfSmH2, rλH2S, where mH2, rdenote the hydrogen reacted per unit mass of input feed and λH2Sis the ratio between the total hydrogen input and the reacted hydrogen. The overall conversion in the guard and hydrotreating processes has been evaluated semi-empirically throughout the so-called transfer coefficients fi, kand fH2, kSdefined, respectively, as the mass distribution of the atomic composition of the dry fraction of the input feed among the different phases produced and the distribution of the hydrogen reacted in the process among phases. From these definitions, the mass flow rate of the oil, gas, and aqueous phases are calculated from
 |
4 |
Similarly, the atomic composition can be evaluated from
 |
5 |
Table 6 shows the values of the transfer coefficients fi, kSand fH2, kestimated using the measurements obtained during pilot-scale tests shown in Table 5.
Table 6. Estimated Values of the Transfer Coefficients for Dry Atomic Composition of the Feed fi, kSand the Reacted Hydrogen fH2, kamong Product Phases in the Overall Guard and Hydrotreating Processes.
| process | guard (overall) | hydrotreating | ||||
|---|---|---|---|---|---|---|
| phase | oil | gas | aqueous | oil | gas | aqueous |
| reacted H2 | 0.45 | 0.33 | 0.22 | 0.70 | 0.21 | 0.09 |
| C (dry feed) | 0.925 | 0.006 | 0.069 | 0.899 | 0.006 | 0.096 |
| H (dry feed) | 0.987 | 0.012 | 0.002 | 0.924 | 0.007 | 0.069 |
| O (dry feed) | 0.498 | 0.028 | 0.474 | 0.656 | 0.005 | 0.338 |
| N (dry feed) | 0.871 | 0.000 | 0.129 | 0.388 | 0.464 | 0.148 |
| S (dry feed) | 0.363 | 0.314 | 0.323 | 0.231 | 0.128 | 0.641 |
| Fe (dry feed) | 0.56 | 0.82 | ||||
Table 5. Yields and Composition Measured for the Guard Reactor and Hydrotreating Reactor during Pilot Tests.
| test | biocrude | guard bed stage 1 | guard bed stage 2 | hydrotreating |
|---|---|---|---|---|
| operational conditions | ||||
| temperature (deg. C) | 290 | 290 | 360 | |
| pressure (bar-a) | 100 | 100 | 100 | |
| WHSV (1/h) | 0.5 | 0.5 | 0.4 | |
| H2 reacted (% wt.) | 0.36 | 0.17 | 1.57 | |
| yields (wt. %) | ||||
| liquid hydrocarbons | 93.28 | 94.55 | 95.34 | |
| gas | 4.20 | 2.80 | 3.81 | |
| water | 2.52 | 2.65 | 0.85 | |
| oil composition (dry basis) | ||||
| carbon (wt. %) | 76.84 | 79.54 | 79.94 | 83.68 |
| hydrogen (wt. %) | 9.15 | 10.41 | 10.73 | 11.90 |
| nitrogen (wt. %) | 3.82 | 3.80 | 3.74 | 2.86 |
| sulfur (wt. %) | 0.76 | 0.42 | 0.31 | 0.14 |
| oxygen (wt. %) | 9.43 | 5.83 | 5.28 | 1.42 |
| iron (ppm) | 1054 | 467 | 84 | <10 ppm |
| gas composition (mol % H2 free) | ||||
| CH4 | 14.39 | 22.84 | 16.74 | |
| ethane | 3.40 | 11.11 | 1.15 | |
| propane | 1.13 | 5.75 | 0.77 | |
| butane | 0.28 | 1.52 | 1.30 | |
| CO2 | 77.70 | 34.78 | 67.83 | |
| NH3 | 0.00 | 0.00 | 1.08 | |
| H2S | 3.10 | 24.01 | 11.14 | |
3.5.2. Hydrocracking
The overall hydrocracking processes have been defined in terms of the total input and reacted mass of hydrogen, denoted, respectively, by mH2HCand mH2, rper unit feed mass, the mass yields of main phase products mkHCwith k denoting liquid oil, gas, and process water. Table 8 shows the values of mH2, mH2, rHC, and mkused in the analysis, which are based on reported data on hydrocracking of heavy distillate derived from distillation of hydrotreated pyrolysis oil.30 To the knowledge of the authors, there are no experimental or simulation results in the literature on hydrocracking of heavy distillates produced from hydrotreating of sewage sludge-derived HTL oils. Therefore, although the composition of heavy distillates derived from pyrolysis oil and HTL oil can differ significantly, the approach of using pyrolysis oil results has been considered here just to obtain estimates of the product yields and the consumption of hydrogen and catalysts. It has been assumed that all the remaining S and N heteroatoms present in the heavy distillation feed to hydrocracking are reduced to H2S and NH3, with mass yields calculated from mH2SHC= (yS, f/MWS)MWH2S and mNH3HC= (yN, f/MWN)MWNH3, where yi, fHCdenotes the atomic composition of the heavy distillate feed to the hydrocracking process.
Table 8. Total Mass of Hydrogen Input and Reacted and Mass Yields of Products per Unit Feed Mass for (1) Processing HTL Biocrude in a Guard Reactor and Hydrotreater Based on Experimental Results and (2) Hydrocracking of Heavy Distillate Derived from Distillation of Hydrotreated Pyrolysis Oil30.
| process | guard reactor and hydrotreating | hydrocracking |
|---|---|---|
| reacted hydrogen (% wt.) | 19.2 | 5.4 |
| light hydrocarbons yield (% wt.) | 1.25 | 11.5 |
| naphtha yield (% wt.) | 13.55 | 26.9 |
| middle distillate yield (% wt.) | 29.9 | 61.6 |
| heavy distillate yield (% wt.) | 55.3 |
3.5.3. Distillation
The distillation fractions considered in the analysis with specification of the boiling temperature range reference values for the carbon numbers, specific gravity, carbon and hydrogen content, average molecular weight, and high heating value are shown in Table 7. The overall distillation of the mixture of the liquid oil products from the hydrotreating and hydrocracking process is defined in terms of the mass flow rates of the input feed and the distillation products, calculated, respectively, from ṀjDIST= ṀO+ ṀOHCand Ṁj= ṀOHTyj+ ṀOHCyj. Here, the subscript j represents each of the distillation fractions, i.e., light hydrocarbon gases, naphtha, middle distillate, and heavy distillate, and yjHTand yjare the mass fractions of j in the hydrotreating and hydrocracking oils shown in Table 8. The values of yjHTare shown in Table 7 calculated from yj= ∫TBj, minTBj, maxfO(TB)dTB, which represents the integral of the measured distillation curve fOHT(TB), shown in Figure 3, over the range of boiling temperatures specified by each distillation fraction, as shown in Table 8. The values of yjin Table 7 are based on literature data.30
Table 7. Definition and Reference Properties of the Distillation Fractions Considered in the Analysis.
| distillation fraction | light hydrocarbons | naphtha | middle distillate | heavy distillate |
| TBP (deg. C) | <80 | 80–210 | 210–340 | <340 |
| carbon number | <C5 | C6–C10 | C11–C20 | >C20 |
| specific gravity | 0.66 | 0.78 | 0.82 | 0.94 |
| carbon (% wt.) | 82.43 | 84. 5 | 86.21 | 87.1 |
| hydrogen (% wt.) | 16.1 | 14.2 | 13.5 | 12.9 |
| molecular weight (g/mol) | 102 | 130 | 200 | 425 |
| HHV (MJ/kg) | 48.5 | 46.7 | 45.8 | 44.3 |
Figure 3.
Measured distillation curve for the oil produced hydrotreating of HTL biocrude during pilot tests by Steeper Energy.
3.5.4. Amine Gas Treatment
The amine system is considered as a package defined in terms of the total amine flow rate required in the absorber V̇MEAABSand the specific heating Q̇h, cooling Q̇cGT, and electricity Ẇelrequired per unit mass of amine injected into the absorber. The mass flow rate of amine to the absorber is calculated from V̇MEAABS= γ[ V̇CO2/χCO2ABS+ V̇H2S/χH2SABS+ V̇NH3/χNH3ABS]/Pg, where χjABSare the solubility of CO2, H2S, and NH3 in MEA at the absorber temperature and γ is the ratio of the actual flow ratio of amine relative the flow rate required to reach equilibrium. Here, we have considered a linear dependence for the solubility. The total heating, cooling, and electric power demands by the overall amine system have been assumed to be proportional to the flow rate of amine to the absorber and are calculated from [Q̇h, Q̇cGT, Ẇel] = V̇MEAABS[qh, qcGT, wel], where qhGT, qc, welGTare the specific heating, cooling, and electricity required per unit volume of amine, assumed to be constant and equal to,31 respectively, 0.27 MJ/kg, 0.40 MJ/kg, and 46.5 kWh/kg.
3.6. Material and Energy Balances for the Overall Conversion of Biofuels from Sewage Sludge
The main material and energy flows for the production and upgrading of HTL biocrude from sewage sludge are shown, respectively, in Tables 9 and 10. The conversion of sewage sludge to biocrude exhibits a conversion efficiency of 73.4% on energy basis and 29.4% on dry-mass basis. The overall mass and energy yields of combined naphtha and middle distillate from sewage sludge on dry basis are approximately 19 and 60%, where the naphtha fraction represents about 45% of the total. Losses in the chemical energy during the hydrothermal liquefaction of the sewage sludge are in the form of dissolved organic components in the aqueous phase, short-chain hydrocarbons in the gas phase, and unconverted non-dissolved carbon in the solid residue, which represent about 17, 4.9, and 4.6% on energy basis, respectively. Chemical energy from the biocrude, which are not converted to naphtha and middle distillate, are mainly in the form of light hydrocarbon gases and dissolved organics on the process water after hydrotreating and hydrocracking, which represent about 11.7% and 10% of the biocrude energy. The net heat demand in the production of biocrude represents approximately 20% of the total feedstock energy, of which 8% is covered by combustion of the HTL gas and stripped ammonia and the remaining 12% by an external source of natural gas. The overall heat demand by the complete upgrading process is 4.9% of the chemical energy content in the biocrude, of which 1.2% is used by hydrotreating, 1.3% by hydrocracking, and 2.4% by distillation. The total hydrogen consumed in the overall upgrading is approximately 4.0% wt. relative to the biocrude feed. The distribution of hydrogen consumption between the hydrotreating, including the guard reactor, and the hydrocracking processes is approximately the same. The CO2 emissions by the overall sludge to the biofuel conversion process are about 0.58 kg CO2 per dry kg sewage sludge, of which 64% is biogenic from combustion of HTL gases and light hydrocarbons produced during upgrading and 36% is fossil-based from combustion of natural gas for covering the heat demand of the biocrude production and from reforming of natural gas for production of the make-up hydrogen used in the upgrading.
Table 9. Main Material Flows Based on an Input Feed of 1 Dry Ton Sewage Sludge.
| sewage sludge | ton | 4.29 |
| slurry to HTL | ton | 5.44 |
| HTL oil | ton | 0.29 |
| HTL aqueous phase | ton | 3.96 |
| HTL solid | ton | 1.03 |
| HTL gas | ton | 0.16 |
| base (NaOH) to hydrothermal liquefaction | kg | 14.5 |
| catalyst (K2CO3) to hydrothermal liquefaction | kg | 6.2 |
| citric acid to phase separation | kg | 5.0 |
| MEK to phase separation | kg | |
| combustion air | Nm3 | 879 |
| natural gas | kg | 35.2 |
| lime (gas cleaning) | kg | 4.04 |
| flue gas | ton | 1.33 |
| flue gas | Nm3 | 1593.8 |
| CO2 to air | ton | 0.35 |
| fossil | ton | 0.16 |
| biogenic | ton | 0.19 |
| NH3 to SCR | kg | 0.69 |
| catalyst to SCR | kg | 0.14 |
| dry scrubber residue | kg | 4.04 |
| MVR concentrate bleed to disposal | ton | 0.28 |
| process water from biocrude production | ton | 3.96 |
| emissions to water from biocrude production plant | dm3 | 2.07 |
| light HC from stripping | kg | 14.55 |
| organic liquid to distillation | ton | 0.316 |
| light HC from distillation | kg | 14.6 |
| naphtha range from distillation | kg | 85.8 |
| diesel range from distillation | kg | 100.3 |
| heavy fraction from distillation | kg | 115.2 |
| guard reactor catalyst | kg | 0.046 |
| hydrotreating catalyst | kg | 0.037 |
| hydrocracking catalyst | kg | 0.018 |
| process water | dm3 | 0.031 |
| sour gas from separation | kg | 30.19 |
| total make-up H2 | kg | 11.87 |
| consumption in guard reactor | kg | 1.52 |
| consumption in hydrotreating | kg | 4.13 |
| consumption in hydrocracking | kg | 6.22 |
| total fuel gas consumption | kg | 8.59 |
| to fire heater before hydrotreating | kg | 3.25 |
| to fire heater before distillation column | kg | 1.72 |
| to fire heater before hydrocracking | kg | 3.62 |
| make-up amine consumption | kg | 0.11 |
| fresh water | dm3 | 2.06 |
| GHG emissions | kg CO2.eq | 576.5 |
| from combustion of natural gas | kg CO2.eq | 157.6 |
| from combustion of HTL gases | kg CO2.eq | 188.8 |
| from make-up H2 production | kg CO2.eq | 83.7 |
| from light HC combustion (fired heaters) | kg CO2.eq | 146.4 |
| solid residue | kg | 0.10 |
Table 10. Main Energy Flows (MW) Based on an Input Feedstock Chemical Energy of 1 MW Based on HHV.
| chemical energy raw sludge | 1 |
| slurry to HTL | 1.17 |
| chemical energy oil product | 0.734 |
| chemical energy aq. effluent after phase separation | 0.385 |
| chemical energy gas after phase separation | 0.049 |
| chemical energy solid residue after phase separation | 0.046 |
| chemical energy MVR concentrate bleed to disposal | 0.066 |
| chemical energy treated water to disposal | 0.01 |
| heating slurry preparation | 0.103 |
| heating slurry to HTL | 0.457 |
| heat recovery from HTL product cooling | 0.418 |
| heating of HTL process water before MVR | 0.009 |
| MVR condensate cooling | 0.044 |
| natural gas consumption | 0.12 |
| heat recovery boiler | 0.126 |
| heat recovery after SCR | 0.034 |
| heat lost from flue gas to air | 0.011 |
| chemical energy biocrude feed | 0.734 |
| chemical energy light gases | 0.074 |
| chemical energy naphtha | 0.281 |
| chemical energy middle distillate | 0.321 |
| chemical energy H2 consumed | 0.027 |
| heating biocrude before hydrotreating | 0.009 |
| heating organic liquid before distillation | 0.017 |
| heating organic liquid before hydrocracking | 0.010 |
4. Equipment Scale-Up and Cost Analysis
The economic performance of the overall production and upgrading of the biocrude has been evaluated as a function of the dry sludge feed capacity to the biocrude production and the biocrude feed capacity to the upgrading, denoted respectevely by ṀSHTLand ṀBC. This analysis involves scale-up of the equipment included in the process flow diagrams shown in Figures 1 and 2 using parametric models for the equipment design and costs described bellow. It has been assumed that the feedstock composition and all the process design parameters shown in Tables 1 and 2 are constant for the whole range of plant capacity and that the variations of all mass and enthalpy flow rates have a linear dependency on the conversion capacity. The capacity range used for the biocrude upgrading is 7 to 70 ton/day, which corresponds to the biocrude produced from sewage sludge with a feed capacity of 30–300 dry ton/day based on a biocrude conversion efficiency of approximately 29.4%.
4.1. Equipment Design Models
The size
of silos and liquid storage tanks are specified by their volume V, calculated from V = ṀftRφ/(1 –
ϕ), as a function of the feed mass flow rate Ṁf and density ρf and the residence time tR. The parameters ϕ and φ denote, respectively, the porosity
of the bed material and the empty volume to material filled volume
ratio. For all storage equipment, the value of φ is constant
and equal to 1.2. Conveyors are defined in terms of the length L and diameter D. The diameter is calculated
from D = [(Ṁf/ρf)/vf]1/2 based on the characteristic velocity
of the feed along the conveyor vf, assumed
to be constant and equal to 5 cm/s. The length of the conveyor has
been estimated considering the layout distance between equipment.
The total electric driving power is calculated from Ẇel = (ẆN + ẆM + ẆH)/ηm, where ẆN = DL/20 is the empty power loss, ẆM = fmgṀfL is the load power due to friction losses caused by the weight of
the material, and ẆH = gṀfH is the power required to overcome an elevation difference H. Here, the constant fm is
a progress resistance representing an artificial friction coefficient
for the moving material, g is the gravitational constant,
and ηm is the electric to mechanical
power efficiency of the motor assumed to be constant and equal to
0.9. Pumps and fans are specified by the total electric power, calculated
from Ẇel = ṀfgHp/ηm, where ηm is the electric to mechanical power efficiency of the motor and Hp = (ΔP + ΔPp + ΔPeq)/ρg + dz represents the total head. Here, ΔP is the pressure increase required by the main downstream
process, ΔPp and ΔPeq are the pressure losses in the piping and auxiliary
equipment between the pump or fan and the downstream process equipment,
respectively, and dz is the difference in elevation.
This formulation assumes that the variations of the kinetic energy
due to reduction or increase in flow velocities are negligible. Also,
the effect of the elevation in the pressure drop when calculating
fans has been neglected. Piping losses have been calculated from ΔPp = (Kp + fpLp/Dp)ρv2/2, where v, ρ, Lp, and Dp denotes, respectively, the internal fluid
velocity, the fluid density, the pipe length, and the pipe diameter.
The parameters fp and Kp represent the friction coefficient for the fully developed
internal flow and the pressure drop coefficients due to elbows, valves,
and fittings. The friction coefficient fp is calculated as a function of the Reynolds number Re = ρvDp/μ and the pipe roughness
number ε, using Poiseuille’s law fp = 64/Re for laminar flows (Re < 2300) and the Colebrook–White correlation32
, with 
for turbulent flows (Re < 4000). The characteristic piping length is a function
of the
equipment layout and thus the plant capacity. Compressors have been
specified based on the total electric power calculated from Ẇel = (Ṁg/ρg)Pi[(Po/Pi)(1 – 1/k) – 1][k/(k – 1)]/ηm, where Ṁg and ρg are the mass flow rate and density of the gas, Pi and Po are the
inlet and discharge pressure, ηm is the electric
to mechanical power efficiency of the motor, and k is the polytropic coefficient assumed to be constant and equal to
0.8. Specification of the heat exchangers has been evaluated based
in the total duty Q̇th and heat transfer area Ath. Calculations
of the duty depends on whether the heat exchanger is used as a cooler
or heater. For a heater, the duty Q̇th = Ṁhcp, h(Thin– Th) is a function of the inlet and outlet
temperatures, the mass flow rate, and the specific heat capacity of
the hot stream. Likewise, the duty for a cooler is calculated from Q̇th = Ṁccp, c(Tcin– Tc), where the subscript denotes
here the cold stream. Based on the duty, the total heat transfer area
is calculated from Ath = Q̇th/(UthLMTD), where LMTD is the log mean temperature difference
and Uth is the overall heat transfer coefficient.
It has been assumed that all heat exchangers are designed as a shell
and tube in counterflow, with the overall heat transfer coefficient
calculated from Uth = {dt/Nutkt + (dt/2kt)Ln[dt/(dt –
2t)] + (dt – 2t)/Nusks}−1. In this equation, Nut, kt, Nus, and ks are the Nusselt number and the thermal conductivity
for the flows in the tubes and shell, respectively, and dt, t, and kt are the tube diameter, thickness, and thermal conductivity. Assuming
horizontal staggered tubes, the Nusselt numbers for the shell can
be estimated from 
. The Nusselt number for fluids inside the
tubes are calculated using correlation for internal flows in cylindrical
tubes NudtB= 0.023RetPr0.3. Stirred tanks are specified by the total volume and dimensions
of the tank and the total electric power of the impeller. The tank
volume is calculated from VT = (Ṁf/ρf)tR in terms of the residence
time tR and the feed mass flow rate Ṁf and density ρf. All stirred tanks are assumed to be cylindrical
with the diameter and height calculated from D =
(4VT/πkHD)1/2 and L = kHD(4VT/πkHD)1/2, where kHD is the height
to diameter ratio, which is assumed to be constant and equal to 2.
The power consumption is calculated from ẆelSR= NPρNIDI5/ηel, where Nl is the impeller rotational speed (rpm), DI is the impeller diameter, and NP = 346.7/ReI + 1.27 is the
impeller power number based on the Reynolds number ReI = (NI/60)DI2ρf/μf. The filter press is modeled after the plate and frame press
type with a constant filtration rate and an applied pressure differential ΔP = (V̇/Ac)tcηαc, where αc = 1/κc is the filter cake resistivity, κc is the permeability, and η is the
viscosity. The filter cake resistivity αc = f(dh) is a
function of the hydraulic pore diameter dh = 4ϕc/[(1 – ϕc)SV] of the cake
through the Kozeny–Carman Equation αc ∼ 5(1 – ϕc)2/(ϕcSV)2. Here, SV is the inner solid
surface area per unit volume, which for spherical particles of an
average diameter (ds) is equal to SV = 6/ds. A centrifuge
is described in terms of the total electric power, calculated from Ẇel = kw(Ṁf/ρf), where kw is
the specific electricity consumption per unit volume of the input
feed, assumed to be constant and equal to 1.4 kWh/m3. The
size of all catalytic reactors is specified by their volume V, calculated as a function of the feed mass flow rate Ṁf and the weight hourly
space velocity WHSVcatspecified for
the catalyst from V = (ṀfK/WHSVcat)ρcat–1φ/(1 – ϕ), where ρcat is the density of the catalyst and the parameters ϕ
and φ denote, respectively, the porosity of the bed material
and the empty volume to material filled volume ratio. For all catalytic
reactors, the values of ϕ and φ are constant and equal
to 0.15 and 1.4, respectively.
4.2. Electric Power Loads
The calculated values of the electric power load for the biocrude production and upgrading as a function of the feed capacity are shown in Tables 11 and 12. The electric power consumption for all the main systems involved in the overall conversion process behaves almost linearly with the feed capacity, indicating that the effect of the higher pressure lost with larger plant layouts is small in the overall power consumption. The highest contributions to the power consumption in the biocrude production plant correspond to the MVR unit, the slurry pump, and the centrifuges in the phase separation systems, accounting approximately to 55, 26, and 15% of the total. The total electric load for the overall upgrading process is small, about 2%, compared to the biocrude production, the main contributions corresponding to the feeding pumps to the hydrotreating and hydrocracking processes, the amine system, and the hydrogen compressor, which account for 35.9, 14.7, 12.5, and 10% of the total.
Table 11. Electric Loads (kW) as a Function of the Plant Capacity for the Overall Production of Biocrude from Sewage Sludge.
| sewage sludge capacity (dry-ton/day) | 30 | 50 | 100 | 150 | 200 | 250 | 300 |
| slurry prep. and HTL | 85.74 | 144.17 | 293.78 | 447.64 | 605.20 | 766.11 | 766.11 |
| sludge pump | 0.550 | 0.807 | 1.365 | 1.862 | 2.324 | 2.761 | 2.761 |
| feedstock conveyor | 1.354 | 2.349 | 5.090 | 8.110 | 11.351 | 14.778 | 14.778 |
| slurry preparation stirred tank | 1.947 | 4.558 | 14.456 | 28.402 | 45.866 | 66.519 | 66.519 |
| catalyst conveyor | 0.041 | 0.064 | 0.115 | 0.163 | 0.209 | 0.254 | 0.254 |
| base conveyor | 0.068 | 0.106 | 0.192 | 0.273 | 0.351 | 0.428 | 0.428 |
| slurry pump | 81.776 | 136.288 | 272.561 | 408.832 | 545.101 | 681.368 | 681.368 |
| phase separation | 57.29 | 95.48 | 190.97 | 286.45 | 381.93 | 477.41 | 477.41 |
| first centrifuge | 32.446 | 54.076 | 108.152 | 162.228 | 216.304 | 270.380 | 270.380 |
| mixing vessel after centrifuge | 4.563 | 7.605 | 15.211 | 22.816 | 30.421 | 38.027 | 38.027 |
| second centrifuge | 20.281 | 33.802 | 67.603 | 101.405 | 135.206 | 169.008 | 169.008 |
| gas treatment | 14.19 | 23.03 | 44.97 | 66.67 | 88.21 | 109.64 | 109.64 |
| combustion air compressor | 1.078 | 1.749 | 3.420 | 5.072 | 6.713 | 8.346 | 8.346 |
| natural gas supply | 0.174 | 0.276 | 0.523 | 0.762 | 0.997 | 1.230 | 1.230 |
| exhaust fan | 5.013 | 7.892 | 15.038 | 21.995 | 28.844 | 35.619 | 35.619 |
| SCR package | 2.814 | 4.690 | 9.381 | 14.071 | 18.761 | 23.451 | 23.451 |
| dry scrubber package | 4.573 | 7.622 | 15.243 | 22.865 | 30.487 | 38.109 | 38.109 |
| lime conveyor | 0.097 | 0.149 | 0.266 | 0.374 | 0.476 | 0.575 | 0.575 |
| filter dust conveyor | 0.100 | 0.154 | 0.275 | 0.387 | 0.493 | 0.595 | 0.595 |
| thermal fluid pump | 0.339 | 0.503 | 0.828 | 1.141 | 1.435 | 1.715 | 1.715 |
| water treatment | 200.04 | 333.32 | 666.45 | 999.54 | 1332.61 | 1665.66 | 1665.66 |
| process water pump | 1.058 | 1.738 | 3.417 | 5.082 | 6.737 | 8.387 | 8.387 |
| MVR package | 198.658 | 331.096 | 662.193 | 993.289 | 1324.385 | 1655.482 | 1655.482 |
| CW pump condensate cooler | 0.092 | 0.134 | 0.223 | 0.302 | 0.374 | 0.442 | 0.442 |
| condensate pump | 0.208 | 0.317 | 0.550 | 0.777 | 0.994 | 1.206 | 1.206 |
| concentrate pump | 0.025 | 0.037 | 0.065 | 0.091 | 0.116 | 0.140 | 0.140 |
Table 12. Electric Loads (kW) as a Function of the Plant Capacity for the Overall Production of Naphtha and Middle Distillate from Biocrude.
| sewage sludge capacity (dry-ton/day) | 30 | 50 | 100 | 150 | 200 | 250 | 300 |
| hydrotreating | 1.978 | 3.236 | 6.342 | 9.413 | 12.479 | 15.523 | 18.558 |
| biocrude pump | 1.402 | 2.332 | 4.655 | 6.977 | 9.311 | 11.636 | 13.961 |
| natural gas compressor | 0.043 | 0.064 | 0.115 | 0.160 | 0.202 | 0.243 | 0.282 |
| combustion air fan | 0.210 | 0.333 | 0.629 | 0.915 | 1.195 | 1.472 | 1.746 |
| exhaust fan | 0.323 | 0.507 | 0.943 | 1.362 | 1.770 | 2.172 | 2.569 |
| separation and fractionation | 0.059 | 0.098 | 0.197 | 0.295 | 0.394 | 0.492 | 0.590 |
| naphtha column | 0.030 | 0.049 | 0.098 | 0.148 | 0.197 | 0.246 | 0.295 |
| diesel column | 0.030 | 0.049 | 0.098 | 0.148 | 0.197 | 0.246 | 0.295 |
| hydrocracking | 0.985 | 1.624 | 3.209 | 4.785 | 6.357 | 7.926 | 9.492 |
| heavy distillate pump | 0.574 | 0.951 | 1.894 | 2.836 | 3.777 | 4.718 | 5.659 |
| fuel gas compressor | 0.020 | 0.032 | 0.059 | 0.085 | 0.111 | 0.136 | 0.161 |
| combustion air fan | 0.160 | 0.263 | 0.518 | 0.771 | 1.023 | 1.274 | 1.524 |
| exhaust fan | 0.232 | 0.378 | 0.738 | 1.094 | 1.447 | 1.798 | 2.147 |
| gas treatment and hydrogen recycle | 0.854 | 1.351 | 2.511 | 3.646 | 4.729 | 5.820 | 6.898 |
| amine system | 0.487 | 0.811 | 1.623 | 2.434 | 3.246 | 4.057 | 4.869 |
| H2-rich recycle compressor | 0.221 | 0.325 | 0.528 | 0.722 | 0.903 | 1.074 | 1.239 |
| make-up H2 compressor | 0.146 | 0.214 | 0.361 | 0.490 | 0.580 | 0.688 | 0.791 |
4.3. Capital Cost
The capital cost has been evaluated in terms of the total permanent investment CTPI, calculated from
 |
6 |
The first term in this formula represents the sum of the purchase and installation cost of each equipment included in the process design shown in Figures 1 and 2, calculated from CPI, k = CP, kB(Sk/Sk)nk(I/IB)finst, k, where CP, kBand Skare the base-case equipment purchase cost and equipment size, Sk is the actual size of equipment, nk is the equipment scale factor, finst, k is the equipment installation factor, and I/IB is the price index ratio between the actual year and the reference year where the base case purchase cost function evaluated based on the Chemical Engineering Plant Cost Index (CEPCI). Table 13 lists the values for the equipment cost parameters used in the analysis. All the plant costs are updated to 2021. The parameters fi in eq 6 are additional capital cost factors associated with land, civil work for site preparation and construction of buildings, engineering and contingencies for civil work and process equipment, and project development and licenses. Representative values33,36 for fi are listed in Table 14.
Table 13. Parameters for Calculating the Purchase and Installation Cost for the Equipment.
| equipment | base specification | base purchase cost | installation factor | scale factor | base year | ref | |
|---|---|---|---|---|---|---|---|
| screw conveyor | 33.5 | t/h | 0.350 | 2.10 | 0.80 | 2002 | (34) |
| belt conveyor | 28.6 | t/h | 0.070 | 2.10 | 0.80 | 2009 | (34) |
| sludge pump | 45.0 | kWe | 0.175 | 2.47 | 0.70 | 2017 | (35) |
| sludge tank | 76.7 | m3 | 0.174 | 2.47 | 0.70 | 2010 | (35) |
| shell & tube heat exchanger | 7.8 | MW | 0.080 | 1.24 | 0.60 | 2010 | (34) |
| solid storage silo | 2821.0 | kg/h | 0.055 | 2.10 | 0.70 | 2013 | (34) |
| HTL slurry pump | 333.0 | kW | 0.470 | 2.84 | 0.80 | 2011 | (37) |
| HTL slurry preheater | 17.5 | MW | 1.970 | 2.72 | 0.70 | 2012 | (37) |
| HTL reactor | 5.4 | m3 | 0.270 | 2.47 | 1.00 | 2013 | (37) |
| product cooler | 74.9 | MW | 5.540 | 2.72 | 0.70 | 2013 | (37) |
| gas separator | 12.0 | l/s | 0.190 | 2.73 | 0.84 | 2007 | (34) |
| flash tank | 20.0 | m3 | 0.014 | 2.73 | 0.71 | 2007 | (34) |
| condenser | 0.2 | MW | 0.067 | 1.67 | 0.53 | 2017 | (34) |
| centrifuge | 500 | kg/h | 0.149 | 2.47 | 0.38 | 2007 | (34) |
| gas burner | 1.0 | MW | 0.250 | 1.79 | 0.74 | 2007 | (34) |
| syngas injector | 0.0 | kg/s | 0.008 | 1.24 | 0.32 | 2007 | (34) |
| natural gas burner | 0.5 | MW | 0.002 | 3.03 | 0.16 | 2007 | (34) |
| gas compressor | 15.0 | kW | 0.014 | 2.47 | 0.70 | 2017 | (34) |
| natural gas storage tank | 20.0 | m3 | 0.014 | 2.97 | 0.71 | 2007 | (34) |
| exhaust fan | 109.2 | kW | 0.600 | 2.47 | 0.70 | 2008 | (34) |
| SCR (package) | 53362.5 | Nm3/h | 0.833 | 2.47 | 0.70 | 2001 | (37) |
| water pump | 3.7 | kWe | 0.009 | 2.84 | 0.80 | 2013 | (34) |
| scrubber + bag filter | 53362.5 | Nm3/h | 2.054 | 2.47 | 0.70 | 2001 | (37) |
| cylindrical atmospheric tank | 1.5 | m3 | 0.017 | 2.73 | 0.93 | 2007 | (34) |
| cooler | 0.4 | MW | 0.060 | 1.67 | 0.53 | 2017 | (34) |
| MVR package | 0.4 | m3/h | 0.490 | 2.47 | 0.50 | 2018 | (27) |
| acid storage tank | 1981 | kg/h | 0.100 | 1.73 | 0.71 | 2020 | (34) |
| acid pump | 1981.0 | kg/h | 0.023 | 2.47 | 0.70 | 2010 | (34) |
| mixing vessel with agitator | 500 | kg/h | 0.016 | 2.22 | 0.70 | 2019 | (34) |
| oil tank | 76.7 | m3 | 0.174 | 2.47 | 0.70 | 2010 | (30) |
| oil feed pump | 8.5 | l/s | 0.104 | 4.42 | 0,8 | 2013 | (30) |
| guard reactor | 20 | m3 | 0.231 | 2.73 | 0,81 | 2010 | (30) |
| hydrotreating reactor | 24,948 | kg/h | 6.05 | 1.77 | 1 | 2013 | (30) |
| HP separator (3-phase) | 9.15 | kg/s | 0.44 | 1.24 | 0.84 | 2007 | (30) |
| LP stripper | 0.924 | kg/s | 0.065 | 2.10 | 0.8 | 2017 | (30) |
| naphtha column | 6.62 | kg/s | 0.49 | 2.08 | 0.7 | 2013 | (30) |
| naphtha condenser | 4.08 | kg/s | 0.031 | 5.66 | 0.7 | 2013 | (30) |
| naphtha accumulator | 4.08 | kg/s | 0.125 | 4.95 | 0.7 | 2013 | (30) |
| naphtha reflux pump | 4.08 | kg/s | 0.032 | 5.32 | 0.8 | 2013 | (30) |
| naphtha column reboiler | 2.94 | kg/s | 0.048 | 3.49 | 0.7 | 2013 | (30) |
| diesel column | 2.94 | kg/s | 0.30 | 2.36 | 0.7 | 2013 | (30) |
| diesel condenser | 1.49 | kg/s | 0.013 | 8.23 | 0.7 | 2013 | (30) |
| diesel accumulator | 1.49 | kg/s | 0.031 | 5.36 | 0.7 | 2013 | (30) |
| diesel reflux pump | 1.49 | kg/s | 0.007 | 4.90 | 0.8 | 2013 | (30) |
| diesel column reboiler | 1.11 | kg/s | 0.027 | 3.54 | 0.7 | 2013 | (30) |
| heavy distillate pump | 1.1 | kWe | 0.046 | 2.47 | 0.7 | 2017 | (30) |
| hydrocracking reactor | 10,886 | kg/h | 2.62 | 2.43 | 1 | 2013 | (30) |
| sour gas cooler | 0.4 | MW | 0.060 | 1.67 | 0.53 | 2017 | (30) |
| amine (package) | 41.9 | kg/s | 5.45 | 2.69 | 0.65 | 2005 | (31) |
| hydrogen compressor | 15.0 | kW | 0.014 | 2.47 | 0.70 | 2017 | (30) |
| fired heater | 0.2 | MW | 0.143 | 1.88 | 0.6 | 2013 | (30) |
Table 14. Average Unit Prices for Consumables and Utilities and Financial Assumptions Considered for Calculating the Operating Costs and the Minimum Fuel Selling Price.
| direct operational cost | unit cost |
|---|---|
| gate fee sludge treatment, €/ton | 250 |
| sulfuric acid, €/ton | 62 |
| citric acid, €/ton | 640 |
| NaOH, €/ton | 300 |
| hydrothermal liquefaction catalyst (K2CO3), €/ton | 1400 |
| enzyme (protease), €/ton | 1240 |
| MEK, €/ton | 1440 |
| lime, €/ton | 120 |
| MgO, €/ton | 150 |
| NH3, €/liter | 44.8 |
| catalyst guard reactor, €/liter | 31.4 |
| catalyst hydrotreating, €/liter | 31.4 |
| catalyst hydrocracking, €/liter | 31.4 |
| solid residue disposal, €/ton | 40 |
| process water disposal, €/m3 | 8.3 |
| fresh water, €/m3 | 0.5 |
| electricity, €/kWh | 1.0 |
| natural gas, €/MWh | 24 |
| H2 production cost (SMR), €/kg | 3.75 |
| CO2 emissions (fossil), €/ton | 25 |
| amine (MEA), €/kg | 2.9 |
| labor average annual income, k€/year | |
| managers | 162 |
| O&M manager | 88 |
| engineers | 96 |
| maintenance technician | 59 |
| shift supervisor | 66 |
| shift operators | 59 |
| administration | 37 |
| site and building maintenance | 37 |
| overhead factor (operators only), % | 20 |
| labor overhead charge rate fraction | 1.25 |
| administration cost, % total permanent investment | 2 |
| insurance cost, % of the total permanent investment | 1 |
| loan interest rate, % | 7 |
| return of investment, % | 10 |
| equity to debt ratio | 30/70 |
| plant lifetime, years | 25 |
| construction time, years | 2 |
| commissioning time, years | 1 |
| annual operating time, h | 8000 |
Figure 4 shows the variation with the feed capacity of the total specific installed equipment cost per unit mass flow rate of dry feed of the complete biocrude production plant and for the main systems involved in the feedstock to biocrude conversion. The specific installed equipment cost for the biocrude production plant varies between 0.44 and 0.23 M€/dry-ton/day for plant capacities between 30 and 300 dry-ton/day. The largest contributions to the total installed equipment cost for the baseline design are the HTL and MVR units, representing approximately 70% of the total. As the plant capacity increases, the contribution from the MVR to the total installed cost also reduces due to its lower scale factor. The installed equipment cost for the overall biocrude production have a dependency with the plant capacity, which corresponds to an average scale factor of 0.7. Figure 5 shows the specific equipment installed cost per unit biocrude mass flow rate as a function of the feed capacity for the complete upgrading process and for the main systems. The total cost of equipment required for the upgrading varies between 130 and 80 k€/ton/day for the biocrude feed capacity range considered. The main contribution to the total equipment cost corresponds to the guard and hydrotreating system, representing approximately 44% of the total, followed by the hydrocracking and fractionation systems, which account for 25 and 26%. The cost of the equipment for cleaning and recycling the hydrogen-rich sour gas is relatively low, representing 5% of the total equipment cost. The installed equipment cost for the complete biocrude upgrading exhibits a power dependency with the feed capacity corresponding to an average scale factor of 0.78. The calculated values of the total permanent investment with the contribution of the different cost factors for the biocrude production and the biocrude upgrading are shown in Tables 15 and 16. All the costs associated to the development and construction of the biocrude production and the biocrude upgrading plants represent about 65% of the total permanent investment, approximately 80% of which is due to civil work, engineering, and contingencies.
Figure 4.
Variation of the specific installed equipment cost per unit mass flow rate of sewage sludge feed as a function of the biocrude production capacity: (left) total; (right) distribution among main systems, i.e., slurry preparation and HTL, phase separation, HTL water treatment, and HTL gas treatment.
Figure 5.
Variation of the specific installed equipment cost per unit mass flow rate of biocrude feed as a function of the biocrude upgrading capacity: (left) total; (right) distribution among main systems, i.e., hydrotreating (including the guard reactor), separation and fractionation, hydrocracking, and sour gas treatment with H2 recirculation.
Table 15. Annual Operating Cost and Income (M€/Year) for the Biocrude Production as a Function of the Feed Capacity.
| 30 | 50 | 100 | 150 | 200 | 250 | 300 | |
|---|---|---|---|---|---|---|---|
| total permanent investment | 29.90 | 42.26 | 68.54 | 91.67 | 113.07 | 133.33 | 152.76 |
| equipment installed cost | 13.33 | 18.81 | 30.45 | 40.67 | 50.12 | 59.05 | 67.61 |
| slurry preparation and HTL | 5.41 | 8.25 | 14.72 | 20.76 | 26.55 | 32.17 | 37.66 |
| phase separation | 1.64 | 2.22 | 3.36 | 4.31 | 5.15 | 5.93 | 6.65 |
| gas treatment | 1.68 | 2.40 | 3.89 | 5.17 | 6.32 | 7.39 | 8.40 |
| water treatment | 4.59 | 5.95 | 8.48 | 10.43 | 12.09 | 13.56 | 14.89 |
| chemicals (initial batch) | 0.22 | 0.36 | 0.72 | 1.09 | 1.45 | 1.82 | 2.18 |
| piping | 0.87 | 1.22 | 1.98 | 2.64 | 3.26 | 3.84 | 4.39 |
| electrical system | 0.67 | 0.94 | 1.52 | 2.03 | 2.51 | 2.95 | 3.38 |
| instrumentation & control system | 0.60 | 0.85 | 1.37 | 1.83 | 2.26 | 2.66 | 3.04 |
| project costs | 14.22 | 20.08 | 32.50 | 43.40 | 53.48 | 63.01 | 72.15 |
| land | 1.33 | 1.88 | 3.05 | 4.07 | 5.01 | 5.91 | 6.76 |
| site preparation | 0.73 | 1.03 | 1.67 | 2.24 | 2.76 | 3.25 | 3.72 |
| foundation and buildings | 2.67 | 3.76 | 6.09 | 8.13 | 10.02 | 11.81 | 13.52 |
| plant engineering | 2.71 | 3.82 | 6.19 | 8.27 | 10.19 | 12.00 | 13.74 |
| contingency | 3.61 | 5.10 | 8.25 | 11.02 | 13.58 | 16.00 | 18.32 |
| project development and licenses | 0.73 | 1.03 | 1.67 | 2.23 | 2.75 | 3.24 | 3.71 |
| commissioning | 2.44 | 3.44 | 5.57 | 7.44 | 9.17 | 10.80 | 12.37 |
| annual operating cost (M€) | 2.56 | 3.85 | 6.89 | 9.81 | 12.67 | 15.48 | 18.27 |
| consumables and utilities | 0.867 | 1.445 | 2.895 | 4.350 | 5.809 | 7.271 | 8.738 |
| base (NaOH) to HTL | 0.044 | 0.073 | 0.146 | 0.219 | 0.292 | 0.365 | 0.438 |
| catalyst (K2CO3) to HTL | 0.087 | 0.146 | 0.291 | 0.437 | 0.583 | 0.728 | 0.874 |
| acid to phase separation | 0.032 | 0.054 | 0.108 | 0.162 | 0.216 | 0.270 | 0.324 |
| MEK (phase separation) | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 |
| natural gas | 0.268 | 0.447 | 0.895 | 1.342 | 1.789 | 2.237 | 2.684 |
| lime (gas cleaning) | 0.005 | 0.008 | 0.016 | 0.024 | 0.033 | 0.041 | 0.049 |
| NH4OH (25% NH3) to SCR | 0.001 | 0.002 | 0.003 | 0.005 | 0.006 | 0.008 | 0.009 |
| fresh water | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 |
| electricity | 0.429 | 0.715 | 1.435 | 2.160 | 2.890 | 3.623 | 4.359 |
| labor | 0.20 | 0.21 | 0.24 | 0.27 | 0.29 | 0.31 | 0.33 |
| maintenance | 0.27 | 0.38 | 0.61 | 0.81 | 1.00 | 1.18 | 1.35 |
| insurance and taxes | 0.60 | 0.85 | 1.37 | 1.83 | 2.26 | 2.67 | 3.06 |
| administration and services | 0.30 | 0.42 | 0.69 | 0.92 | 1.13 | 1.33 | 1.53 |
| emissions to air | 0.040 | 0.066 | 0.132 | 0.198 | 0.264 | 0.330 | 0.396 |
| emissions to water | 0.174 | 0.290 | 0.579 | 0.869 | 1.158 | 1.448 | 1.737 |
| disposal of solid residue | 0.113 | 0.188 | 0.377 | 0.565 | 0.753 | 0.942 | 1.130 |
| income (M€) | 3.12 | 5.20 | 10.40 | 15.60 | 20.80 | 26.00 | 31.20 |
| sludge | 3.12 | 5.20 | 10.40 | 15.60 | 20.80 | 26.00 | 31.20 |
Table 16. Annual Operating Cost and Income (M€/Year) as a Function of the Biocrude Feed Capacity.
| biocrude feed capacity (ton/day) | 8.8 | 14.7 | 29.4 | 44 | 58.7 | 73.4 | 88.1 |
| total permanent investment (M$) | 2.60 | 3.43 | 5.85 | 8.05 | 10.11 | 12.08 | 13.98 |
| equipment installed cost | 1.16 | 1.71 | 2.91 | 3.99 | 5.01 | 5.98 | 6.93 |
| hydrotreating | 0.47 | 0.69 | 1.21 | 1.68 | 2.13 | 2.57 | 2.99 |
| phase separation and fractionation | 0.24 | 0.35 | 0.60 | 0.82 | 1.02 | 1.21 | 1.40 |
| hydrocracking | 0.37 | 0.54 | 0.92 | 1.26 | 1.58 | 1.89 | 2.18 |
| sour gas treatment and hydrogen recycle | 0.08 | 0.11 | 0.18 | 0.23 | 0.27 | 0.32 | 0.36 |
| chemicals (initial batch) | 0.02 | 0.03 | 0.06 | 0.08 | 0.11 | 0.14 | 0.17 |
| piping | 0.08 | 0.11 | 0.19 | 0.26 | 0.33 | 0.39 | 0.45 |
| electrical system | 0.06 | 0.09 | 0.15 | 0.20 | 0.25 | 0.30 | 0.35 |
| instrumentation & control system | 0.05 | 0.08 | 0.13 | 0.18 | 0.23 | 0.27 | 0.31 |
| plant development costs | 1.24 | 1.42 | 2.43 | 3.33 | 4.18 | 5.00 | 5.78 |
| land | 0.12 | 0.17 | 0.29 | 0.40 | 0.50 | 0.60 | 0.69 |
| site preparation | 0.06 | 0.09 | 0.16 | 0.22 | 0.28 | 0.33 | 0.38 |
| foundation and buildings | 0.23 | 0.34 | 0.58 | 0.80 | 1.00 | 1.20 | 1.39 |
| plant engineering | 0.24 | 0.26 | 0.44 | 0.60 | 0.75 | 0.90 | 1.04 |
| contingency | 0.31 | 0.34 | 0.58 | 0.80 | 1.00 | 1.20 | 1.39 |
| project development and licenses | 0.06 | 0.05 | 0.09 | 0.12 | 0.15 | 0.18 | 0.21 |
| commissioning | 0.21 | 0.17 | 0.29 | 0.40 | 0.50 | 0.60 | 0.69 |
| total annual operating cost | 4.48 | 5.84 | 8.42 | 10.43 | 12.13 | 13.63 | 14.97 |
| consumables and utilities | 4.150 | 5.432 | 7.814 | 9.635 | 11.153 | 12.469 | 13.639 |
| biocrude | 3.695 | 4.673 | 6.297 | 7.361 | 8.120 | 8.679 | 9.091 |
| catalyst to guard reactor | 0.029 | 0.048 | 0.096 | 0.144 | 0.192 | 0.240 | 0.288 |
| catalyst to hydrotreating | 0.023 | 0.039 | 0.078 | 0.117 | 0.156 | 0.195 | 0.235 |
| catalyst to hydrocracking | 0.011 | 0.019 | 0.037 | 0.056 | 0.074 | 0.093 | 0.112 |
| make-up hydrogen | 0.385 | 0.642 | 1.284 | 1.926 | 2.568 | 3.210 | 3.851 |
| amine | 0.003 | 0.005 | 0.011 | 0.016 | 0.022 | 0.027 | 0.032 |
| fresh water | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 | 0.000 |
| electricity | 0.003 | 0.005 | 0.010 | 0.015 | 0.020 | 0.025 | 0.030 |
| labor | 0.17 | 0.17 | 0.17 | 0.17 | 0.17 | 0.17 | 0.17 |
| maintenance | 0.02 | 0.03 | 0.06 | 0.08 | 0.10 | 0.12 | 0.14 |
| insurance and taxes | 0.05 | 0.07 | 0.12 | 0.16 | 0.20 | 0.24 | 0.28 |
| administration and services | 0.03 | 0.03 | 0.06 | 0.08 | 0.10 | 0.12 | 0.14 |
| emissions to air | 0.058 | 0.097 | 0.194 | 0.290 | 0.387 | 0.484 | 0.581 |
| emissions to water | 0.003 | 0.004 | 0.009 | 0.013 | 0.017 | 0.021 | 0.026 |
| income for selling of light gases | 0.07 | 0.11 | 0.22 | 0.33 | 0.44 | 0.55 | 0.66 |
4.4. Operating Cost and Income
Tables 14 and 15 show the calculated results of the capital investment, annual operating cost, and annual income as a function of the feed capacity for the biocrude production and the biocrude upgrading, respectively. The annual operating costs are calculated as the sum of variable direct operational cost Cop, d dependent on the annual processing of feedstock, the fixed indirect operational costs Cop, i required for having the plant in activity, and the maintenance costs Cmaint. The direct operational costs include the purchase of consumables and utilities and the cost of the emissions to air and the disposal of solid residues and effluents. For the guard, hydrotreating, and hydrocracking reactors, the annual consumption of catalyst per unit hourly flow rate of the feed to the reactor is calculated in terms of the weight hourly space velocity of the catalyst from ṀcatS= (1/WHSVcat)(tp/tcatS) with tp and tcatdenoting the annual production time and the life time of the catalyst (h). These costs are calculated based on the individual rates of consumption or production obtained from the mass and energy flows reported in Tables 8 and 9 with unit prices listed in Table 12 assuming an annual operating time of 8000 h. The annual indirect operational costs labor, administration, and insurance. The total annual labor cost has been evaluated from Clabor = ∑jNj[cr, j(1 + flb) + fOH, jcOH] , where the subscript j denotes the personnel categories, and Nj, cr, j, flb, fOH, j, and cOH represent the annual man-hour of personnel required, the hourly rate, the labor burden factor, the overhead factor, and the overhead cost factor, respectively. The personnel categories and the values for the labor cost used are shown in Table 13. The number of personnel required has been estimated based on individual main systems proportional to the purchase and installation costs, except for management, which is assumed to be constant. The costs for administration and insurance are evaluated as a percentage of the total permanent investment according to Table 13. The specific operating list for the biocrude production plant varies between 0.23 and 0.17 k€/dry-ton for feed capacities of sewage sludge in the range 30 to 300 dry-ton/day. Here, the cost of consumables and utilities represents approximately 25% of the total operating cost, with the electricity and the natural gas consumption contributing by 49 and 31%, respectively. The annual income from treating sewage sludge, considering a typical gate fee of 250 €/ton, can cover the total annual operating cost. Considering the biocrude upgrading in a centralized refinery, the specific operating cost varies between 1.65 and 0.54 k€/ton for biocrude feed capacities in the range of 8.8 and 88 ton/day, where about 84% of the total is due to the cost of the biocrude feed. The second main contribution is the cost of producing the make-up hydrogen from natural gas at the refinery, which represents approximately 4%. The annual cost of catalyst replacement, assuming a unit price of 31.4 €/liter and an operational lifetime of two years, represents only 1.1% to the total operating cost.
4.5. Levelized Cost of Biocrude and Minimum Fuel Selling Price
The levelized costs of biocrude and the minimum fuel selling price (MFSP), denoted by cbc and cbf, are defined as the average prices for the biocrude and biofuels, respectively, per unit energy produced so that the overall net present value (NPV) for the total permanent investment and over its lifetime becomes zero. Based on these definitions, cbc and cbf are calculated using the formulas
 |
7 |
and
 |
8 |
Here, r is the expected return of investment, CTPI, iK, COP, i, and CREV, iKare the annual distributions of the annual total investment, operating costs, and revenues over the plant lifetime, with K = PROD, UPG denoting the biocrude production plant and the biocrude upgrading plant, respectively. In this notation, ϵbc is the annual energy efficiency of the sewage sludge to biocrude conversion, ϵD and ϵN are the annual energy efficiency of the biocrude to diesel and naphtha conversion, pN is the market price of naphtha relative to diesel, and tp, i is the annual production time assumed to be 8000 h. The financial assumptions used in eqs 7 and 8 and the unit prices for calculation of revenues are shown in Table 14. Figure 6 shows the variation as a function of the sewage sludge feed capacity of the biocrude production cost and the minimum fuel selling price.
Figure 6.
Variation as a function of the sewage sludge feed capacity of the levelized cost of biocrude production cost per liter of biocrude production (dashed line) and the minimum fuel selling price per liter (solid line) considering full investment in a new stand-alone upgrading unit and use of existing upgrading equipment at refinery.
The results for the MFSP shown in Figure 6 have considered two different scenarios, i.e., full investment in a new stand-alone upgrading unit and use of existing upgrading equipment at refinery. In this second scenario, it is assumed that all equipment required for the upgrading of the biocrude are available at refinery and no capital investment is required, except for the initial batch of consumables, and only the annual operating costs and revenues are used for evaluating the minimum fuel selling price. The levelized cost of biocrude, which exhibits a monotonic decrease with the feed capacity, is in the range of 36.8 to 8.9 €2021/GJ per unit energy and 1.4 to 0.34 €2021/liter per unit volume for plant capacities between 30 and 300 dry-ton/day. These results are in agreement with the values found in the literature.20,22 If all the capital investment in a new upgrading plant is included, the average MFSP varies between 64.7 and 21.5 €2021/GJ per unit total energy of diesel and naphtha and 2.4 and 0.8 €2021/liter per unit volume of diesel equivalent for the range of the sewage sludge feed capacity used in the analysis. The parameters that impact the most on the MFSP are the biocrude price, the production cost of the make-up hydrogen, and the annual expenditure due to capital investment. Assuming that all equipment needed for upgrading of the HTL biocrude is available and can be utilized at the refinery with full replacement of fossil-derived feeds with HTL biocrude, the MFSP can be reduced only by approximately 6–7% for the production capacities considered. The results for the MFSP obtained from this analysis are also in line with the latest values obtained by Snowden-Swan et al.20
5. Conclusions
Production of liquid biofuels for road transportation can be achieved by (decentralized) direct conversion of the sewage sludge to an intermediate oil phase, so-called biocrude, via hydrothermal liquefaction at near-critical water conditions and further upgrading of the biocrude to naphtha and middle distillate at a centralized conventional refinery. The gas product from liquefaction can be co-combusted with natural gas for production of the net heat demand by the overall biocrude production process, which represents approximately 12% of the chemical energy contained in the sewage sludge. The aqueous effluent from liquefaction can be treated by air stripping for separation of the dissolved ammonia, which is combusted with the HTL gas, followed by mechanical vapor recompression. The overall mass and energy yields of biocrude are approximately 29.4 and 73.4% of the sewage sludge, respectively, with MBSP varying between 36.8 and 8.9 €2021/GJ per unit energy and between 1.4 and 0.34 €2021/liter per unit volume for sewage sludge feed capacities in the range of 30–300 dry-tons/day. The main costs contributing to the MBSP are the purchase and installation of the HTL process water treatment systems and capital cost associated to engineering and construction of the biocrude production plant. The overall upgrading process includes multi-stage catalytic hydrotreating of the biocrude for reduction of inorganics and S, N, and O heteroatoms, separation of the sour gas and water from the liquid oil, fractionation of the hydrotreated oil by distillation, and catalytic hydrocracking of the heavy distillate separated from fractionation. The main distillation products are naphtha and middle distillate, which represent gasoline and diesel pools, respectively, at the refinery. The overall mass and energy yields of combined naphtha and middle distillate from sewage sludge on dry basis is approximately 19 and 60%, where the naphtha fraction represents about 45% of the total. When considering investment in a new stand-alone unit for upgrading the biocrude, the minimum fuel selling price that can be achieved varies between 64.7 and 21.5 €2021/GJ per unit total energy of diesel and naphtha and 2.4 and 0.8 €2021/liter per unit volume of diesel equivalent for biocrude feed capacities in the range of 8.8 to 88 ton/day. The main contribution to the overall MFSP comes from the cost of biocrude, which represents about 60%, followed by the purchase and installation cost of process equipment. If existing equipment at refinery can be used for upgrading the biocrude, thus avoiding capital cost due to equipment and plant development and construction, the minimum fuel selling price reduces by 7%. Sewage sludge is considered a model urban waste feedstock posing the main challenges for the conversion to biofuels, i.e., high nitrogen and metal contents. Therefore, the overall production costs reported in this document are expected to be higher than those for biofuels produced from biocrude derived from other urbane waste fractions with lower contents of metals, N, S, and O.
Acknowledgments
This project has received funding from the European Union’s Horizon 2020 Research and Innovation 415 Programme under Grant Agreement No. 818413. (NextGenRoadFuel - Sustainable Drop-In Transport fuels 416 from Hydrothermal Liquefaction of Low Value Urban Feedstocks). The authors acknowledge the contribution from Steen Brummerstedt Iversen, Claus Uhrenholt Jensen, and Julie Katerine Rodriguez from Steeper Energy to this work, particularly in the revision of the process design and the values on chemical consumption by the HTL process and the upgrading of the HTL biocrude.
The authors declare no competing financial interest.
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