Energy Integration of a Hydrotreatment Process for ... - ACS Publications

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Energy Integration of a Hydrotreatment Process for Sustainable Biojet Fuel Production Claudia Gutiérrez-Antonio,*,† Araceli Guadalupe Romero-Izquierdo,‡ Fernando Israel Gómez-Castro,‡ and Salvador Hernández‡ †

Facultad de Química, Universidad Autónoma de Querétaro, Cerro de las Campanas s/n Col. Las Campanas, Querétaro, Querétaro, 76010, México ‡ Departamento de Ingeniería Química, División de Ciencias Naturales y Exactas, Campus Guanajuato, Universidad de Guanajuato, Noria Alta s/n, Guanajuato, Guanajuato, 38010, México S Supporting Information *

ABSTRACT: Biojet fuel has been identified as the most promising alternative for sustainable development of the aviation sector. There are several processes for biojet fuel production, being the hydrotreating of vegetable oils the most attractive. Considering that one of the reactions in the process is exothermic, it is possible to use this energy to reduce the heating requirements and the environmental impact of the process. Therefore, in this work we propose the energy integration of the hydrotreating process to produce biojet fuel, considering Jatropha curcas as renewable raw material. Results show a decreasing in heating−cooling services when energy integration is performed in the process; however, the investment costs are increased due to the equipment required to perform the integration. Therefore, the total annual costs of processes, with and without energy integration, are similar. Nevertheless, energy integration allows a reduction of 86% in the environmental impact of the process. The application of energy integration allows reducing the carbon footprint and increasing the sustainability of renewable aviation fuel. contributes 3%.4 In spite of these small percentages (5% combined), the forecast indicates that, by 2050, the aviation sector (domestic, international, and shipping) will contribute 10−32% of the total CO2 emissions.4 Thereby, in the actual conditions the growth of the aviation sector is not sustainable. In order to have a sustainable growth, the aviation sector has established ambitious objectives such as a 50% reduction in CO2 emissions by 2050, relative to 2005 emissions levels, and, also, neutral growth in carbon dioxide emissions from 2020.5

1. INTRODUCTION Nowadays, the aviation sector faces great challenges. On one side, the growth of this sector is forecast at a rate of 4.8% per year until 2036;1 this finding is consistent with the reports of the International Air Transport Association (IATA), which estimates that air traffic will double by 2020, compared to 2005.2 In order to keep the growth rate, a major amount of fuels is going to be required; however, the depletion in the production of petroleum will not allow having the required jet fuel for the operation of the aircrafts. On the other side, the increase in the amount of used fuel will lead to higher CO2 emissions. According to the International Civil Aviation Organization (ICAO), the passenger transport accounts for 2% of the global CO2 emissions,3 while the International Maritime Organization (IMO) estimates that goods transport © 2016 American Chemical Society

Received: Revised: Accepted: Published: 8165

April 14, 2016 June 9, 2016 July 6, 2016 July 6, 2016 DOI: 10.1021/acs.iecr.6b01439 Ind. Eng. Chem. Res. 2016, 55, 8165−8175

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Industrial & Engineering Chemistry Research

Figure 1. Block diagram of the hydrotreating process.

as camelina, jatropha, and algae oils.12−14 Several demonstration flights have been realized,15 proving that the biojet fuel produced through hydrotreating of vegetable oils reaches the specifications of fossil jet fuel; also, the biojet fuel produced from UOP Honeywell’s process was approved for commercial passenger-bearing flights by ASTM International on July 1, 2011.16 The hydrotreating process considers the transformation of vegetable oil to generate hydrocarbons through two consecutive reactors, then obtaining the desired fractions through distillation, Figure 1. It is important to mention that there is not so much information available for the kinetic model or type of catalyst involved in the hydrotreating process of UOP Honeywell, due to confidentiality issues.2 Therefore, the reported works in the literature related with the complete process are focused on modeling both reactive and separation zones; there are some works related with the experimental hydrotreating of the oils,7,17−20 but only few described the kinetic equations. The hydrotreating process allows producing biojet fuel with the required properties; however, if biojet fuel is expected to be sustainable, it is important to improve the process reducing its energy requirements and its carbon footprint. Therefore, the intensification (through thermally coupled distillation, reactive distillation, or even reactive thermally coupled distillation) and also the energy integration (since the first reaction is highly exothermic) are two main opportunity areas that could help to increase the sustainability of the hydrotreating process. The intensification of the hydrotreating process through thermally coupled distillation sequences has been already proposed;2 the sequences analyzed were direct and indirect thermally coupled distillation sequences, Petlyuk sequence and the dividing wall column. The composition of the effluent of the reaction system was obtained from a previous work of Gutiérrez-Antonio et al.,21 where the modeling of the reactive zone was estimated. The energy consumption and total number of stages of intensified sequences were compared with the indirect conventional sequence, the best of the conventional ones. The results show that it is possible to reduce by 21% the energy consumption of the separation, when intensified sequences are used. On the other side, an interesting opportunity to reduce energy consumption in the hydrotreating process lies in energy integration, using the released energy by the hydrodeoxygenation reaction. This could help to reduce significantly the energy consumption and the carbon footprint of the process. To the authors’ knowledge, this has not been reported in the literature.

These objectives will be reached through the implementation of a four pillar strategy,6 which includes the following: (1) technological improvements in engines and aircraft structures (2) operational improvements by online optimization of flight paths (3) market-based measures, basically emissions trading (4) development of alternative fuels From the previous strategies, the development of alternative fuels has been identified by IATA as one of the most promising ways to significantly reduce the greenhouse gas emissions and achieve independence in fuels. Alternative fuels for aviation, also known as biojet fuel, renewable aviation fuel or synthetic paraffinic kerosene (SPK), are constituted of renewable hydrocarbons in the boiling range of fossil jet fuel. The composition of biojet fuel is very similar to that of the fossil jet fuel, and due to this, its properties are practically the same; however, they do not contain aromatic compounds, whose absence can cause wear in certain types of engines.7 Synthetic paraffinic kerosene can be generated from coal, natural gas, or biomass.7 From the previous alternatives, biomass is the only renewable, and it includes the following feedstock: triglyceride, lignocellulosic, sugar, and starchy. Depending on each type of raw material there are five possible routes to transform biomass into jet fuel: hydrogenated esters and fatty acids (HEFA), Fischer−Tropsch based on biomass (FT), renewable synthesized isoparaffins (SIP), alcohol to jet (ATJ), and hydrogenated pyrolysis oils (HPO). The HEFA, FT, and SIP routes are certified by the ASTM D7566 standard.8 From the certified routes, the HEFA process, or hydrotreating process, is a very interesting alternative to produce biojet fuel. The hydrotreating process was developed by UOP Honeywell,9 and it considers the transformation of triglyceride feedstock through hydrodeoxygenation, hydroisomerizing, and hydrocracking to generate renewable hydrocarbons.10 This process is very similar to the existing ones in the petrorefineries, so in a future scenario where petroleum no longer exists, this infrastructure could be used to produce biojet fuel and other biofuels. The required raw materials are triglycerides, from vegetable and animal sources including waste cooking oils. The hydrotreating process has been proved at pilot plant level,11 using vegetable oils as renewable feedstock.8 UOP Honeywell has produced several thousand gallons of renewable jet fuel from a variety of feedstocks, including first-generation oils such as palm and soybean oils, as well as second-generation oils such 8166

DOI: 10.1021/acs.iecr.6b01439 Ind. Eng. Chem. Res. 2016, 55, 8165−8175

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Industrial & Engineering Chemistry Research Therefore, in this work we propose the energy integration of the hydrotreating process to produce biojet fuel, considering jatropha oil as renewable raw material. In order to estimate the energy requirements in the reaction zone, we model the hydrodeoxygenation, hydrocracking, and hydroisomerizing reactions. Once the energy requirements are determined, we propose energetic integration, so the heat released by the reactions can be used in the process, also taking advantage of the heating/cooling capacity of the streams of the process. The separation zone considers two conventional distillation schemes. The results show interesting findings related to the decreasing in energy consumption and carbon footprint when energy integration is used. The article is organized as follows. First, the selection of the raw material along with the modeling of the reactive zone is presented, which allows estimating the composition of the effluent and also the data of process streams, section 2. In section 3, the methodology used to select the most promising conventional distillation sequences is presented, also considering the composition and flow of the effluent of the hydrocracking−hydroisomerizing reactor. Later, in section 4, the employed methodology to perform the energy integration of the process is described, being the main objective the minimization of the use of external auxiliary services. Also, the calculation of total annual costs and CO2 emissions are presented in section 4. Finally, in section 5 the resulting hydrotreating processes, with and without energy integration, are analyzed in terms of energy consumption, total annual costs, and CO2 emissions.

C tgo

= e −k ′ t →

dC tg dt

(2)

Cj =

dCj −k 2 C tgo(e−k ′ t − 1) → = −k 2C tg dt k′

(3)

Cp =

dCp −k4 C tgo(e−k ′ t − 1) → = −k4C tg dt k′

(5)

Hydrodeoxygenation k′ 14.35

k1 0.04

k12 16, 401, 205, 679.0

k2 0.11 Hydrocracking

k3 1.24

k13 33, 040, 947, 364.0

k4 13.25 k23 74.13

Based on the kinetic expressions, we model the hydrodeoxygenation reactor in the process simulator Aspen Plus using a plug flow reactor at constant temperature. The flow of J. curcas oil feeding the reactor is 100 kg/h, while the hydrogen is fed in a ratio of 1500 mL of H2/mL of oil in order to avoid coke formation.9 The operating conditions of the hydrodeoxygenation reactor are 320 °C and 80 bar; we select these conditions since Sharma et al.32 reported that at 320 °C the conversion to heavier products (C15−C18) is highest. Considering the operating conditions of the reactor, it is clear that both reactants must be conditioned before entering the reactor. From the hydrodeoxygenation reactor we obtain as products carbon dioxide, water, and lineal hydrocarbons. Carbon dioxide and water are separated from the hydrocarbon mixture before they enter the hydrocracking−hydroisomerizing reactor. In the second reactor, two reactions are carried out: hydrocracking and hydroisomerization. The hydrocracking reaction is represented with the model proposed by Shayegh et al.33 for the vacuum gas oil (VGO) catalytic cracking, which considers the formation of coke; however, this reaction path is not considered since the hydrogen/oil ratio utilized avoids coke formation. Therefore, the lumped kinetic model employed is presented next: dC VGO = −k12C VGO2 − k13C VGO2 dt dCgasoline dt dCgl dt

= −k12C VGO2 − k 23Cgasoline

= −k13C VGO2 − k 23Cgasoline

(6)

(7)

(8)

where CVGO is C19−C21 paraffins, Cgasoline is C5−C11 paraffins with its isomers, and Cgl is C1−C4 paraffins with their isomers; k12, k13, and k23 are the kinetic constants, which are also shown in Table 1. In the second reactor there is no additional feed of hydrogen, since the amount introduced in the hydrodeoxygenation reactor is enough for the whole reactive section. Based on the kinetic expressions, we model the hydrocracking− hydrosiomerizing reactor in the process simulator Aspen Plus using a plug flow reactor at constant temperature; the operating

(1)

dC L − k1 C tgo(e−k ′ t − 1) → = −k1C tg k′ dt

(4)

Table 1. Kinetic Parameters for the Hydrodeoxygenation and Hydrocracking Reactions

= −k′C tg

CL =

dCj −k 3 C tgo(e−k ′ t − 1) → = −k 3C tg dt k′

where Ctg is triglycerides, CL is light compounds (C5−C8), Cj is middle compounds (C9−C14), Ch is heavy compounds (C15− C18), and Cp is oligomerized compounds (>C18); k′, k1, k2, k3, and k4 are kinetic constants, which are presented in Table 1.

2. MODELING OF REACTIVE ZONE We selected as renewable feedstock Jatropha curcas, which is a second-generation raw material with high productive potential in Mexico according to INIFAP.22 Several state governments in Mexico are encouraging the production of J. curcas, which is a promising option as a second-generation renewable feedstock for the production of biofuels in Mexico, especially biojet fuel. The composition of J. curcas oil, taken from the work of ́ 23 is presented next: triolein (42 wt %), Herrera-Martinez, trilinolein (44 wt %), tripalmitin (11 wt %), and tristearin (3 wt %). As mentioned before, the hydrotreating process consists of the transformation of vegetable oil through hydrogenating, deoxygenating, isomerizing, and selective hydrocracking to generate renewable fuels, which are purified later.10 Thereby, the process is integrated by two consecutive reactors and a conventional distillation sequence. The study of the hydrotreating reaction of vegetable oils is of recent interest; experimental research has been performed for soybean,24,25 rubber seed,26 ucuhuba,26 canola,27 pine,27 sunflower,28 rapeseed,29,30 and cottonseed.31 In particular, the hydrodeoxygenation of jatropha oil was studied by Sharma et al.,32 who determined a lumped kinetic model: C tg

Ch =

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Industrial & Engineering Chemistry Research conditions are 480 °C and 80 bar, since at these conditions the conversion of cracking products is highest. It is important to mention that these conditions are the same for the hydrocracking and hydroisomerization reactions. Also in the hydrocracking−hydroisomerization reactor, the isomerization is modeled with the kinetic model proposed by Calemma et al.:34 dC isoparaffins dt

= k 2Cparaffins

requirements, searching for the minimization of the external auxiliary services. Pinch analysis is one of the best-known methods of process integration and can be used for various objectives, primarily for heat recovery, fresh water minimization, or recovery of organic solvents.35 The methodology to determine the minimum use of external services along with the location of the pinch point is presented next: • identification of temperatures, flow rates, and enthalpies of all process streams for each conventional hydrotreating process • definition of minimum difference in temperature, and adjusting of temperatures of the process streams • definition of temperature intervals, according to the adjusted temperatures, and performing of energy balance in each defined interval • construction of the heat cascade, and construction of the heat exchanger network All the heating/cooling requirements which cannot be fulfilled by the process streams are supplied by steam or cooling water. Heat exchangers are then introduced to perform the integration between the streams. The equipment and utilities costs are computed for both conventional and integrated hydrotreating processes, using the cost equations shown in Turton et al.,36 which are updated with the CEPCI (Chemical Engineering Process Cost Index). The installation prices are also updated by the CEPCI index. The capital cost (purchase plus installation cost) is annualized over a period which is often referred to as the plant lifetime. We considered the plant life as 5 years, being 8400 the number of operating hours per year. Therefore, the total annual cost is calculated with

(9)

where Cisoparaffins and Cparaffins are the concentrations of isoparaffins and paraffins, respectively; k2 is the kinetic constant, with a value of 0.00239 min−1. From the reactive section a renewable hydrocarbon stream is obtained, which includes light components, naphthas, biojet fuel, and green diesel. The separation of each product is realized through distillation, as described in section 3.

3. MODELING OF SEPARATION ZONE The renewable hydrocarbon stream leaves the reactive section at 480 °C and 80 bar, which represent a high pressure condition for a stream that is going to be fed to a distillation train. Therefore, we proposed the use of a turbine to perform the conditioning of the stream before its feeding to the distillation train; also, the turbine will allow generating electrical energy in the process. Once the pressure of the hydrocarbon stream is decreased, it is fed to a distillation train where four products are obtained: light gases (C1−C4), naphthas (C5−C7), biojet fuel (C8−C16), and green diesel (C17−C21). Considering conventional distillation, these products can be separated through five distillation sequences. However, the separation of the light components implies the use of refrigerant as a cooling service; due to this, we decided to eliminate all the sequences where the light gases were not obtained in the first distillation column. Also, with the objective of decreasing the use of refrigerant, we are proposing a partial condenser in the first column of the train. Therefore, we consider only two distillation schemes: direct and direct−indirect. The distillation trains were designed with the shortcut methodology Fenske− Underwood−Gilliland−Kirkbride, through the module DSTW in Aspen Plus. The recoveries of the key components were established in 99%, while the thermodynamic model employed was Braun K10. The obtained designs are simulated in the Radfrac module of Aspen Plus, in order to consider material and energy balances along with the rigorous calculation of phase equilibria. Based on the modeling of the reactive and separation zones, two hydrotreating processes were defined. The modeling of the reactive zone is common to both processes, being the difference between them the use of direct sequence or direct−indirect sequence in the separation zone. From here, we will call conventional hydrotreating processes those where there is no energy integration, and integrated hydrotreating processes those where energy integration is performed. In section 4 the methodology to perform energy integration is described.

annual capital cost = capital cost/plant lifetime

(10)

total annual cost (TAC) = annual operating cost + annual capital cost

(11)

It is important to mention that the operating costs were assumed to be utility costs (steam, cooling water, and refrigerant) along with the electric energy required by the compressors. Finally, the total CO2 emissions include those originated by the electricity used and those due to the generation of steam. The CO2 emissions due to steam generation are calculated with the mass of the fuel (natural gas in this work) along with an emission factor of a natural gas boiler.37 On the other hand, the CO2 emissions due to electricity are computed considering an emission factor associated with the generation of electricity, in Mexico, along with the electric energy consumption in the equipment.

5. ANALYSIS OF RESULTS In this section we will discuss the obtained results. First, we will analyze the conventional hydrotreating processes, including the reactive and separation zones. After that, we will propose two integrated hydrotreating processes, based on the pinch analysis and the information on the process streams. All the processes will be compared in terms of total annual costs, operation cost, equipment costs, and CO2 emissions. Therefore, we will begin with the identification of each analyzed process in terms of scenarios:

4. ENERGY INTEGRATION The conventional hydrotreating processes were simulated in Aspen Plus, in order to know the enthalpy, temperature, and flow of all the process streams involved. After that, we performed a pinch analysis in order to determine the best way to use the available energy in the process according to the 8168


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Figure 2. Scenario 1. Conventional hydrotreating process that includes the direct distillation sequence in the separation zone.

Figure 3. Scenario 2. Conventional hydrotreating process that includes the direct−indirect distillation sequence in the separation zone.

Figure 4. Distribution of lineal chain hydrocarbons in the hydrodeoxygenation reactor.

• scenario 2: conventional hydrotreating process that includes the direct−indirect distillation sequence in the separation zone

• scenario 1: conventional hydrotreating process that includes the direct distillation sequence in the separation zone 8169


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Figure 5. Distribution of hydrocarbons in the hydrocracking−hydroisomerizing reactor.

• scenario 3: integrated hydrotreating process that includes the direct distillation sequence in the separation zone • scenario 4: integrated hydrotreating process that includes the direct−indirect distillation sequence in the separation zone 5.1. Conventional Hydrotreating Processes. Scenarios 1 and 2 are shown in Figures 2 and 3, respectively. From Figures 2 and 3, it can be observed that the reactants must be conditioned in order to be fed to the reactive section of the process. The hydrogen is commercially available at 10 bar and −234.58 °C,38 while the J. curcas oil is considered available at 1 bar and 25 °C. Both reactants must be pressurized and heated in order to reach 80 bar and 320 °C, which are the operating conditions of the hydrodeoxygenation reactor (reactor 1). The specifications of the equipment required to perform the conditioning of 6.6576 kmol/h hydrogen and 0.1144 kmol/h J. curcas oil are presented as Supporting Information (Table S1). Nevertheless, it is important to mention that the main energetic consumption is due to the conditioning of hydrogen. Once the reactants are at the required conditions, they are fed to the hydrodeoxygenation reactor. As was mentioned before, reactor 1 is modeled with the RPlug module at constant temperature. In the hydrodeoxygenation reactor the transformation of J. curcas oil is realized to generate carbon monoxide, carbon dioxide, water, and lineal chain hydrocarbons. The lineal chain hydrocarbons are separated from the other gases in order to be fed to reactor 2. It is important to mention that the hydrocarbon stream also contains hydrogen, since this reactant was fed in excess in the hydrodeoxygenation reactor. This hydrogen is going to be used in the hydrocracking/hydro-

isomerizing reactor; therefore, no additional hydrogen flow is required in this reactor. The distribution of the lineal chain hydrocarbons is shown in Figure 4, where we observe that the main fraction of the stream corresponds to the oligomerized hydrocarbons (>C18). This distribution is consistent with the results reported by Sharma et al.,32 and it is a desirable result since the oligomerized hydrocarbons will be cracked and isomerized in the second reactor. The hydrodeoxygenation is an exothermic reaction, which released 46.05 kW of thermal energy. The hydrocarbon stream that leaves the hydrodeoxygenation reactor is fed to the hydrocracking−hydroisomerizing reactor, which operates at 80 bar and 480 °C. Figure 5 shows the composition profiles in this reactor; we can perceive the presence of light gases (blue tones), naphthas (yellow tones), biojet fuel (red tones), and green diesel (green tones). Also, it is clear that green diesel is the major product generated instead of biojet fuel. The flow rates of each component in these four products are presented in Table 2. From Table 2 we note that the global conversion of the J. curcas oil to hydrocarbons is 85.5%, while the biojet fuel yield is 18.6%. The conversion to biojet fuel is small, but it is consistent with the maximum conversion of 36% reported in the patent of UOP Honeywell.9 Once the hydrocarbons are produced, the hydrogen in excess is removed from the reactor. It is worth mentioning that the hydrogen stream can be recycled to the first reactor; however, this possibility is not explored in this work. The hydrocracking−hydroisomerizing process is endothermic; therefore, 13.03 kW of energy is required. The renewable hydrocarbon stream contains four products: light gases, naphthas, biojet fuel, and green diesel. These products need to be separated in order to commercialize them. 8170


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Industrial & Engineering Chemistry Research Table 2. Products Generated in the Effluent of the Hydrocracking−Hydroisomerizing Reactor product

component

mass flow (kg/h)

light gases

C1 C2 C3 C4 iC5 C5 iC6 C6 iC7 C7 iC8 C8 iC9 C9 C10 C11 iC12 C12 C13 iC16 C14 C15 C16 C17

0.3459 2.5987 4.9011 6.4624 0.2508 12.6710 0.0005 0.0263 0.0786 3.8951 0.0065 0.0012 0.0022 0.1080 0.0041 0.1344 0.1409 6.8319 0.1585 0.0094 8.7107 2.0593 0.4511 2.3313

0.0035 0.0268 0.0506 0.0668 0.0025 0.1310 5.73 × 0.0002 0.0008 0.4027 6.81 × 1.26 × 2.35 × 0.0011 4.28 × 0.0013 0.0014 0.0706 0.0016 9.79 × 0.0900 0.0212 0.0046 0.0241

C18 C19 C20 C21

2.3204 0.1684 0.1083 30.7235

0.0239 0.0017 0.0011 0.3177

naphthas

biojet fuel

green diesel

mass fraction

Table 3. Design Variables of Direct and Direct−Indirect Conventional Distillation Trains

total mass flow (kg/h)

Scenario 1. Conventional Direct Sequence column B1

14.308

16.923 10−6

10−5 10−5 10−5

column B2

no. of stages 17 40 feed stage no. 9 21 type of condenser partial steam total operating press. (bar) 1.01 1.35 distillate flow (kmol/h) 0.3292 0.2197 reboiler duty (kW) 4.09 11.24 Scenario 2. Conventional Direct−Indirect Sequence

18.619

10−5

column B3 65 33 total 1.70 0.0997 4.85

compressor

column B1

column B2

column B3

no. of stages feed stage no. type of condenser operating press. (bar) distillate flow (kmol/h) reboiler duty (kW)

17 9 partial steam 1.01 0.4428 4.09

45 24 total 1.35 0.1233 15.80

40 20 total 1.01 0.0997 4.97

releasing 46.05 kW while the hydrocracking−hydroisomerizing reactor requires 13.03 kW. Both conventional distillation trains performed the required separation: however, the direct scheme consumes 18.8% less energy than the direct−indirect train. At this point all the information regarding the process streams is known; then the pinch analysis can be realized in order to define the integrated hydrotreating processes. 5.2. Integrated Hydrotreating Processes. The reactive and the separation zones of scenarios 3 and 4 are the same as those presented in scenarios 1 and 2, respectively. The difference is the use of the available energy in the process to reduce the external auxiliary services. It is worth mentioning that for the pinch analysis the minimum difference in temperature was set at 10 °C. According to the results presented in section 5.1, the hydrodeoxygenation reactor released 46.05 kW, which is transferred to pressurized water to generate high pressure steam. Thus, the proposed heat exchanger network for scenario 3 is shown in Figure 6, while the detailed information on each heat exchanger of the network is presented as Supporting Information in Table S2. The high pressure steam flow is used to satisfy the reboiler requirements in column B2 along with the thermal necessities in the conditioning of both reactants. The steam flow that leaves the reboiler of column B2 is also used to satisfy the energy requirements of the hydrocracking− hydroisomerizing reactor. Moreover, the resulting medium pressure steam that leaves heat exchanger 3 passes through a reducing valve (v-1) and it is reutilized to supply energy requirements in column B1. However, the overheated steam required in the reboiler of column B3 cannot be generated with energy integration; therefore it has to be produced in a boiler. Here we proposed that saturated vapor is brought and then it is conditioned through a compressor, 2, to be overheated steam. Therefore, the complete scenario 3 is presented in Figure 7, where the energy integration is clearly identified. Scenario 4 has a heat exchanger network similar to the one presented for scenario 3. The hydrodeoxygenation reactor released 46.05 kW, which is transferred to pressurized water to generate medium pressure steam. In this scenario we decided to generate medium pressure steam based on the energy requirements in the distillation train, which are different from those in scenario 3. Thus, the proposed heat exchanger network is shown in Figure 8, while the information on each heat

10−5

35.652

Due to the wide differences in the boiling points of the key components, distillation is selected as the separation process. It is important to remember that the reaction effluent conditions are elevated (80 bar and 480 °C); then the pressure of that stream must be reduced in order to also decrease the operation pressure of the distillation column, favoring the separation and avoiding hazardous conditions in the column. For this, a turbine is proposed along with a cooler to decrease also the temperature of the stream; these units allow the stream to reach a pressure near the atmospheric. The energy consumptions of the equipment are also shown as Supporting Information (Table S1); as expected, a high amount of energy needs to be removed from this stream. Once the hydrocarbon stream is conditioned, it is separated through distillation: conventional direct sequence (scenario 1) and conventional direct−indirect sequence (scenario 2). These schemes are designed considering a recovery of 99% of the key components. Table 3 shows the main parameters of each distillation column for both scenarios. Considering that the reactive zone is common to both scenarios, from Table 3, it is clear that scenario 1 requires less energy consumption, since the distillation train requires 18.8% less energy than the train of scenario 2. Moreover, the total number of stages is very similar; therefore, the associated capital costs are expected to be similar. In all the cases the specified recoveries for the key components are reached. Summarizing, the analysis of the conventional hydrotreating process allows knowing that the conversion of J. curcas oil to biojet fuel is 18.6%, and the hydrodeoxygenation reactor is 8171


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Figure 6. Heat exchange network for scenario 3.

Figure 7. Scenario 3. Integrated hydrotreating process that includes the direct distillation sequence in the separation zone.

Figure 8. Heat exchanger network for scenario 4.

exchanger of the network is presented as Supporting Information in Table S3. In this scenario, the medium pressure steam generated is used to proportionate energy in the reboiler

of column B3 and also in the hydrocracking−hydroisomerizing reactor; moreover, it satisfies the thermal necessities in the conditioning of both reactants. Also, the medium pressure 8172


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Figure 9. Scenario 4. Integrated hydrotreating process that includes the direct−indirect distillation sequence in the separation zone.

Table 4. Summary of Costs (USD/year) of the Four Defined Scenarios scenario

cooling−heating service costs

equipment costs

raw material cost

electricity costs

total annual cost

1 2 3 4

45,520.96 42,941.20 32,091.69 36,949.11

234,017.41 222,317.84 250,199.31 235,632.98

4,470,116.83 4,470,116.83 4,470,116.83 4,470,116.83

2,087.39 2,087.39 2,870.46 2,457.80

4,751,742.59 4,737,463.26 4,755,278.29 4,745,156.73

the savings in operation costs are almost compensated with the increasing in equipment cost. Nevertheless, it is important to observe that investment and operation costs represent only 5.8% of the total annual costs, being the main contributor the cost of the raw material, J. curcas in our case. According to these results, the exploration of other renewable raw materials of minor cost, such lignocellulosic biomass or waste oil or fats, is important. This will help to reduce significantly the total annual costs of the process. If we analyze the CO2 emissions for the four scenarios, we find significant differences. Scenario 3 is the one with the minimum environmental impact, since the CO2 emissions are reduced 86.2% with respect to the scenario without energy integration. The process considered in scenario 3 is the most promising, especially when sustainability certifications are desirable, such as the certification provided by the Roundtable on Sustainable Biomaterials.39 Considering total annual costs along with CO2 emissions, scenario 3 is the best option, since its environmental impact is minimum and the total annual costs are just 0.16% greater than the minimum value found in scenario 2.

steam that is integrated in the network is reutilized in heat exchanger 3, and it is conditioned through a reducing valve to be reutilized to supply the energy requirements in column B1. The overheated steam required in column B2 cannot be generated with energy integration; therefore it has to be produced in a boiler. However, the steam that leaves the bottom of column B2 is used in heat exchanger 6. The complete scenario 4 is presented in Figure 9, where the energy integration is clearly identified. Now, the next step is quantifying the effect of energy integration on total annual costs, operation costs, equipment costs, and CO2 emissions. Table 4 shows the costs of processing J. curcas oil to produce renewable hydrocarbons in the different scenarios, while Table 5 presents the CO2 emissions. The effect of energy integration can be observed in the decreasing of cooling−heating services; however, the number of equipment required to perform the interchange is increased and, as a consequence, capital cost is also increased. In general, scenario 2 is the one with the minor total annual cost; however, all the costs are very close between them since Table 5. Summary of CO2 Emissions (kg of CO2/year) for the Four Defined Scenarios scenario

emissions due to steam generation

emissions due to electricity production

total emissions

1 2 3 4

286 458 714.13 342 814 351.49 39 516 609.68 88 912 371.77

8 692.48 11 445.65 4 869.76 6 953.31

286 467 406.61 342 825 797.15 39 521 479.44 88 919 325.08

6. CONCLUDING REMARKS The energy integration of the hydrotreating process to produce biojet fuel from J. curcas oil has been presented. In order to have all the required information to perform the energy integration, the modeling of reactive and separation zones was carried out. We found that the energy released by the hydrodeoxygenation reactor is almost enough to satisfy all the thermal energy requirements in the process. This allows a 8173


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decrease in the cooling−heating services; however, more equipment is required and thereby the investment costs increase. The total annual costs for each scenario are very similar between them, since the savings in operating costs are diminished with the increase in the equipment cost. However, the main effect of energy integration is on the environmental impact, where a decrease of 86.2% in the total CO2 emissions was observed. It is important to note that investment and operation costs represent only 5.8% of the total annual costs, being the main contributor the cost of the raw material, J. curcas in our case. Therefore, it is necessary to explore processes where lignocellulosic biomass or waste oil or fats can be used as raw materials, in order to develop a sustainable and cost competitive production process for biojet fuel.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.6b01439. Tables including energy consumption of the equipment required to perform the conditioning of the process streams, heat exchanger data for the network of scenario 3, and heat exchanger data for the network of scenario 4 (PDF)

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AUTHOR INFORMATION

Corresponding Author

*Tel.: (+52) 442 192 1200 or (+52) 442 329 93 48. E-mail: [email protected] or [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Financial support provided by CONACyT, through Grant 239765, for the development of this project is gratefully acknowledged. Also, A. G. Romero-Izquierdo was benefited with a grant for the realization of her postgrade studies by CONACyT.

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REFERENCES

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