Integrated biomass thermochemical conversion for clean energy production: Process design and economic analysis

Journal of Environmental Chemical Engineering 7 (2019) 103093

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Integrated biomass thermochemical conversion for clean energy production: Process design and economic analysis

T

⁎

Isah Yakub Mohammeda,c, , Yousif Abdalla Abakra, Robert Mokayab a Department of Mechanical, Materials and Manufacturing Engineering, The University of Nottingham Malaysia, Jalan Broga, Semenyih, 43500, Selangor Darul Eshan, Malaysia b School of Chemistry, University of Nottingham, University Park, Nottingham, NG7 2RD, United Kingdom c Department of Chemical Engineering, Abubakar Tafawa Balewa University, P.M.B 0248, Bauchi, Nigeria

A R T I C LE I N FO

A B S T R A C T

Keywords: Napier grass Thermochemical conversion Process modeling Economic analysis

In this study, conversion of Napier grass bagasse into biofuel was carried out through steady state process modelling and simulation using Aspen Plus software. An integrated thermochemical system consisting pyrolysis, gasification and combustion section for pyrolytic oil production and upgrading, non-condensable gas/syngas and bio-char valorisation was developed in a single flowsheet with capacity of 49 kg/h feedstock. Aspen process economic analyser was used to generate both capital and operating costs estimates, and other investment analysis following first quarter 2016 pricing basis. The simulation result showed that about 68 wt% of the feedstock could be processed into pyrolytic oil at 480 °C with corresponding liquid biofuel yield of 5.35 kg/h, which is equivalent to about 11 wt% of the initial feedstock through in-situ hydrodeoxygenation process. The result of economic evaluation of the entire process plant indicated that a total capital cost of $2.13 M and operating cost/year of $0.3 M is required for 25-year economic life of the project with net present value of $15.18 M at 8% internal rate of return. The standalone biomass refining plant is expected to produce fuel at minimum fuel selling price of $5.81/gallon ($1.45/L) gasoline equivalent, mainly driven by the capital cost. This work presents a systematic combination of process units that will be needed to produce advanced cellulosic biofuels in the near future.

1. Introduction Fossil fuels remain the major cost-efficient and reliable source of energy that currently assure the production of required product quality. Statistics in the past decade have shown that more than 25% of the world energy mainly from diesel and gasoline is consumed in the transportation industry [1]. This trend is expected to continue for decades to come despite the environmental challenges such as emission of greenhouse gases emanating from the utilisation of these fossil fuels. This obstacle can only be overcome by developing competitive clean alternatives that would be affordable and guarantee sustainable production processes. Biomass is the only alternative energy source with carbon in its building block that can be converted into liquid fuel [2]. Currently, the use fuel wood for energy production is becoming more expensive and attention has now shifted to the utilisation of low-quality woody and non-woody biomass, such as agricultural crops and forest residues [3]. Lignocellulosic materials especially those from non-food

crops cultivated on lands that are increasingly marginal for more favoured major crops are potential source of sustainable alternative energy. Napier grass is one of the neglected lignocellulosic non-food energy crops that requires less productive efforts, grown on marginal lands with high biomass yield, typically in the range of 25–35 oven dry tonnes per hectare annually, which corresponds to 100 barrels of oil energy equivalent per hectare. Cultivation of Napier grass follows conventional farming practices. It outcompetes weeds, needs very little or no supplementary nutrients and therefore requires lower establishment costs. It can be harvested up to four times within a year with a ratio of energy output to the energy input of around 25:1 [4]. Valorisation of this materials via thermochemical conversion such as pyrolysis to generate clean fuel precursors have been documented in the literature. According to Strezov, et al. [5], liquid biofuel precursors such as phenol, benzene and other small oxygenates constitutes the major components of Napier grass pyrolytic oil. Similar chemical composition of Napier grass pyrolytic oil has also been reported by Lee,

⁎ Corresponding author at: Department of Mechanical, Materials and Manufacturing Engineering, The University of Nottingham Malaysia, Jalan Broga, Semenyih, 43500, Selangor Darul Eshan, Malaysia. E-mail address: [email protected] (I.Y. Mohammed).

https://doi.org/10.1016/j.jece.2019.103093 Received 20 January 2019; Received in revised form 22 March 2019; Accepted 13 April 2019 Available online 15 April 2019 2213-3437/ © 2019 Elsevier Ltd. All rights reserved.

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efficiency of the process, which could be attributed to lack of proper process integration cum possible non-utilisation of other co-products in the system. The result of economic analysis suggests that gasoline and diesel products can be produced from biomass pyrolysis of pine wood and subsequent hydroprocessing of the pyrolytic oil at a product value of £6.25/GGE, about 35% more than the cost price of fossil gasoline/ gallon. This may be linked to high upgrading cost, which was reported to be about 61% of the total capital investment relative to less than 40% recorded for the pyrolysis step. Shemfe, et al. [17] recently reported a techno-economic study of biofuel from pyrolytic oil via zeolite upgrading route with two options of catalyst regeneration systems. The economic performance was found to be slightly high compared to their earlier study without the catalyst regeneration system. A minimum fuel selling price of £7.20/GGE was recorded, which was reported to be sensitive to fuel yield, operating cost and income tax. The authors reported that improving plant capacity could make the minimum fuel selling price more competitive. However, consideration of plant location may have a positive impact on the minimum fuel selling price since the income tax is a government policy and varies from region to region. The results of techno-economic studies of biofuel production via biomass pyrolysis followed by catalytic upgrading available in the literature is highly dependent on the feedstock type and the process technology used. Currently, information or report on the economic viability of valorisation of Napier grass into biofuel through thermochemical conversion route is very limited in literature. The objective of this study was to conduct a technical and economic analysis of Napier grass pyrolysis to pyrolytic oil and catalytic upgrading into biofuels starting with pre-processed Napier grass bagasse with moisture content of about 45 wt%. This study is focused on developing an integrated thermochemical biomass refining process in one process flow that produces different forms of bioenergy using advanced process simulation software, Aspen Plus® V10 with emphasis on environmental and socioeconomic sustainability.

et al. [6]. Recent study by Ansari and Gaikar [7] also shows the composition of pyrolytic oil from Napier grass constitutes hydrocarbons, phenols, aldehydes and other dissolved organics such as alcohol, acid, aldehyde, phenolic compounds, and sugar chemicals in aqueous phase. The pyrolytic oil yield from these studies in addition to other recent reports in the literature indicate that pyrolytic oil from Napier grass is typically in the range of 32–60 wt% depending the type of pyrolysis process, reactor and process variables used [8–10]. Studies on pyrolytic oil upgrading into transportation fuel are being carried out through the application of different unit operations and unit processes. Though, the investigations are still at early stage of research and development. Issues regarding reaction mechanisms, kinetics and catalyst deactivation remain a challenge. This makes it difficult to forecast technological and economic viability of any upgrading process. Recent study reported by [11] has shown that decarbonylation, decarboxylation, cracking reactions constitute the major reaction pathways in catalytic processing of Napier grass pyrolytic oil using mesoporous zeolite. This finding was further evaluated in terms of yield in a separate study and it was reported that catalytic upgrading of pyrolytic oil over mesoporous zeolite could produce about 38 wt% light fraction, 48 wt% middle distillate and 7 wt% bottom product [12]. Similarly, upgrading of Napier pyrolytic oil via catalytic hydrodeoxygenation with in-situ hydrogen generation from alcohol has also been reported recently. The authors identified Hydrodeoxygenation, hydrogenolysis, hydrogenation, dehydration, demethylation, hydrocracking, decarbonylation and decarboxylation as the main upgrading reactions [13]. Despite the progress on the valorisation of Napier grass into liquid biofuel via pyrolysis and subsequent catalytic upgrading, no economic evaluations of the process have been reported. Though, production of liquid biofuels through biomass pyrolysis is yet to be commercialised due to the high level of investment required and lack of competitiveness with fossil fuels. Application of process modelling and simulation tool becomes very important in order to ascertain process performance and economic parameters so as to determine the most suitable pathway. This would reduce commercialisation risk to a greater extent since the production facilities are generally very difficult to modify once they have been implemented. Techno-economic analysis is a valuable research technique for evaluating the technical and economic feasibility of conceptual process designs. Technical and economic comparison of liquid fuel derived from biomass via thermochemical and biochemical conversion pathways has been reported in the literature. According to Anex, et al. [14], the major discrepancies between the thermochemical (pyrolysis and gasification) and biochemical conversion pathways is mainly the total capital investment with pyrolysis having the lowest value while the biochemical conversion route presents the highest total capital investment, which was attributed to the cost of capital and the feedstock. Thilakaratne, et al. [15] reported economic feasibility of hydrocarbon-based biofuels from woody biomass through initial mild catalytic pyrolysis followed by the upgrading of the partially deoxygenated pyrolysis liquid via hydro processing technology with the application of a ChemCAD software. The result shows 17.70 wt% biomass can be converted into transportation-range fuels with an energy efficiency of 39%. They reported that the thermochemical route with subsequent hydroprocessing achieves 61.1gallon gasoline equivalent (GGE) per metric tonnes of biomass, which is higher relative to the biochemical pathways of ethanol production from cellulosic materials. Their economic analysis shows that the feedstock processing was the major contributor to the operating cost while the minimum fuel selling price was greatly affected by the fuel yield. Shemfe, et al. [16] reported techno-economic performance analysis of biofuel production from pine wood biomass pyrolysis and pyrolytic oil upgrading. The process was modelled using Aspen Plus® software with 72 MT/day biomass and hydroprocessing of pyrolytic oil was adopted for the upgrading. Their findings indicate that initial physicochemical properties of the biomass particularly the moisture content had a considerable impact on the overall energy

2. Material and methods The overall methodology used in this study consist of conceptual process design for the conversion pathway, model development and implementation in Aspen Plus®, economic analysis by discounted cash flow method, sensitivity and uncertainty analyses using Monte Carlo simulation as summarised in Fig. 1. The process design consists of five (5) technical sections: pyrolysis, hydroprocessing, gasification, combustion and power generation. In the pyrolysis section, the bagasse is first processed to reduce the moisture significantly and subsequently charged into pyrolysis reactor to produce pyrolytic oil. The energy requirement for the drying operation is generated in the gasification section. The hydroprocessing section coverts the pyrolytic oil in two stage upgrading reactors with in-situ hydrogen generation from ethanol. In the combustion section, syngas from the gasification and

Fig. 1. Simplified Methodology Flow Chart. 2

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Fig. 2. Process flow diagram for the integrated biomass refining via thermochemical conversion.

2.1.1. Pyrolysis section Biomass fast pyrolysis system is adopted in this study for the production of pyrolytic oil in Aspen platform using its enhanced solid modelling tools. Six (6) Aspen model blocks, the fluidised bed (BFD), decomposition reactor (PYDCM), pyrolysis reactor (PYRX), cyclone (PYCYC), heat exchanger (PYCOOL) and pyrolytic oil collector (PYSEP) were used to represent the unit operations and unit process in the pyrolysis section. The fluidised bed represents the first stage of pyrolysis process, dehydration and mimics the fluid dynamics as obtained in actual fluidised bed reactor system. This stage particularly, the dehydration process of pyrolysis is usually not represented in most of the pyrolysis model reported in the literature. Decomposition reactor, fluidised bed and constant stirred tank reactor are used to represent the pyrolysis process in which biomass is first decompose into individual building block without prior dehydration process [16]. Similarly, Peters, et al. [18] represented pyrolysis process with only two virtual reaction stages that accounted for the biomass decomposition into holocellulose and lignin, and volatilisation of the biomass fragments without implementation of the dehydration step in the Aspen Plus model. In this study, the feedstock (NPGBGS) contains nearly 45 wt% moisture is charged in to the fluidised bed and exited as dried pyrolysis feed (DPYFD) at 95 wt% moisture free using combined hot air from the gasification section AIR2-2 (204 °C), and AIR3-2 (165 °C) from the PYCOOL. The comprehensive characteristics of the NPGBGS and the input components in the Aspen Plus is summarised in Tables 1 and 2 respectively. In the PYDCM represented by RYield reactor, the DPYFD is disintegrated into its chemical building blocks such as hemicellulose, cellulose and lignin, which subsequently undergo pyrolysis reaction in the PYRX at 400–500 °C and residence time of 1.75 s following a multiple pyrolysis reaction kinetics reported by Ranzi, et al. [19] and Peters, et al. [18] in Table 3. The reaction kinetics were implemented in the Aspen plus using its power law reaction tools by specifying the appropriate reacting phase and kinetic factors (pre-exponential factork, reaction order-n, activation energy-Ea). All the kinetic factors (obtained from different sources) are converted to similar units using appropriate conversion factors for better representation of the models in the Aspen Plus software. According the reaction scheme, cellulose, hemicellulose and lignin decomposes to produce the respective monomers as intermediate product. Cellulose first decompose to give glucose,

bio-char together with non-condensable gas from the pyrolysis section are used to produced flue gas, which is afterward utilised to generate high pressure steam for power generation. 2.1. Process model development and implementation in aspen The process design for integrated biomass refining via thermochemical conversion was implemented in ASPEN PLUS® V10 software. The process flow diagram was developed in with non-random two-liquid (NRTL) thermodynamic and property data as shown in Fig. 2. The following the list of assumptions were used:

• Plant capacity is 50 kg/h of Napier grass bagasse (45 wt% moisture, particle size 0.4–1.6 mm) • The average composition of Napier grass bagasse is based on the •

preliminary laboratory analysis (The feedstock proxanal, ultanal and sulfanal attributes are presented in Table 1) Napier grass bagasse is made up of cellulose, hemicellulose and lignin.

Table 1 Average characteristics of Napier grass bagasse. Property Proxanal (wt%) Moisture Content Volatile Mattera Ash Contenta Fixed Carbona HHV(MJ/kg) Ultanal and sulfanal (wt%) dry ash free basis Carbon Hydrogen Nitrogen Sulfur Oxygen Structural composition (wt%) Cellulose Hemicellulose Lignin a

Value

45.00 86.26 1.66 16.74 18.74 49.02 5.78 0.15 0.02 45.03 38.75 19.76 26.99

Dry basis. 3

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Table 2 List of component used in the simulation. Component ID

Type

Component Name

Alias

WATER DEXTR-01 D-XYL-01 SUCRO-01 GLYCI-01 ETHYL-01 PROPY-01 C11H1-01 NPGBGS ETHAN-01 SULFU-01 NITRO-01 OXYGE-01 SULFU-02 SULFU-03 NITRI-01 NITRO-02 HYDRO-01 METHA-01 CARBO-01 CARBO-02 CARBO-03 LEVOG-01 GLYOX-01 ACETA-01 ACETO-01 GLYCO-01 5-HYD-01 METHA-02 PHENO-01 P-CUM-01 FORMA-01 ETHYL-03 PYRUV-01 C11H1-03 CELLU HCELL LIG-H LIG-G LIG-S ASH CYCLO-01 N-HEX-01 BENZE-01 TOLUE-01 XYLEN-01 CYCLO-02 1:3-C-01 CYCLO-03 FURFU-01 FURFU-02 METHY-01 N-OCT-01 CYCLO-04 CYCLO-05 PALLA-01 2-MET-01

Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Nonconventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Solid Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Solid Solid Solid Solid Solid Nonconventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Conventional Solid Conventional

WATER DEXTROSE D-XYLOSE SUCROSE GLYCINE C9H10O2 C10H12O3 C11H14O4 NAPIER GRASS BAGASSE ETHANOL SULFUR NITROGEN OXYGEN SULFUR-DIOXIDE SULFUR-TRIOXIDE NITRIC-OXIDE NITROGEN-DIOXIDE HYDROGEN METHANE CARBON-MONOXIDE CARBON-DIOXIDE CARBON-GRAPHITE LEVOGLUCOSAN GLYOXAL ACETALDEHYDE-1-D ACETONE GLYCOL-ALDEHYDE 5-HYDROXYMETHYLFURFURAL METHANOL PHENOL P-CUMYLPHENOL FORMALDEHYDE ETHYLENE PYRUVIC-ALDEHYDE C11H12O4-N6 CELLULOSE HEMICELLULOSE LIGNIN-H LIGNIN-G LIGNIN-S ASH CYCLOPENTANE N-HEXANE BENZENE TOLUENE XYLENE CYCLOHEXANONE 1,3-CYCLOHEXADIENE CYCLOHEXENE FURFURAL FURFURYL-ALCOHOL METHYLCYCLOPENTANE N-OCTANE CYCLOOCTANE CYCLOHEXANE PALLADIUM 2-METHYL-BUTANE

H2O C6H12O6 C5H10O5 C12H22O11 C2H5NO2-D1 C9H10O2 C10H12O3 C11H14O4 C2H6O-2 S N2 O2 O2S O3S NO NO2 H2 CH4 CO CO2 C C6H10O5-N1 C2H2O2 C2H3DO C3H6O-1 C2H4O2-D1 C6H6O3-N5 CH4O C6H6O C15H16O CH2O C2H4 C3H4O2 C11H12O4-N6 C6H10O5 C5H8O4 C9H10O2 C10H12O3 C11H14O4 C5H10-1 C6H14-1 C6H6 C7H8 C8H10-N1 C6H10O C6H8-E1 C6H10-2 C5H4O2 C5H6O2 C6H12-2 C8H18-1 C8H16-D6 C6H12-1 PD C5H12-2

char particles in the pyrolysis vapour are separated in PYCYC and the char free vapour is quenched into pyrolytic oil by passing through the PYCOOL. The pyrolytic oil and non-condensable gas are separated in PYSEP.

water and char. Afterward, the glucose undergoes series of reactions to give product like levoglucosan, glycoaldehyde, glyoxal, acetaldehyde, acetone in addition to gases like hydrogen (H2), methane (CH4), carbon monoxide (CO), carbon dioxide (CO2) and char. This is in good agreement with literature where cellulose pyrolysis have been reported [20,21]. Hemicellulose pyrolysis first release xylose, formaldehyde, H2 and permanent gases (CO and CO2). Xylose further reacted to produce methanol, ethanol, ethylene, CH4, H2, H2O, CO, CO2 and char [12,18,19]. Lignin on the other hand is classified as p-hydroxyphenyl (H), guaiacyl (G) and syringyl (S) lignin represented by their corresponding monomeric units as C9H10O2, C10H12O3 and C11H14O4. Each of these primary lignin undergoes series of reaction to produce phenolic-based compounds as represented in equation 8–14 (Table 3). Solid

2.1.2. Hydroprocessing section Pyrolytic oil is upgraded in the hydroprocessing section through catalytic in-situ hydrogen generation from ethanol over palladium catalyst using two conversion step. The first reaction step is implemented in a constant stirred tank reactor (UPRXN) using appropriate kinetic factors obtained from the literature (Table 4). The process involves reforming of ethanol to produce hydrogen, dehydration of sugars to furfural followed by hydrogenation of furfural to furfurylalcohol. The 4

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Table 3 Pyrolysis reaction models and kinetics (Ranzi et al., 2008; Peters et al., 2017). S/No 1 2 3 4 5 6 7 8 9 10 11 12 13 14

Reaction

k (1/s)

CELLULULOSE > 0.9 DEXTR-01 CELLULULOSE > 5.3301 WATER + 5.505 CARBO-03 DEXTR-01 > 1.1111 LEVOG-01 DEXTR-01 > 0.95 GLYCO-01 + GLYOX-01 + 0.21 ACETA-01 + 0.2 5-HYD-01 + 0.2 ACETO-01 + 0.2 CARBO-02 + 0.1 CARBO-01 + 0.1 WATER + 0.15 METHA-01 + 0.3 CARBO-03 HEMICELLULOSE > 0.88 D-XYL-01 HEMICELLULOSE > 1.4 HYDRO-01 + 0.8 CARBO-02 + CARBO-01 + 2.2 FORMA-01 D-XYL-01 > 0.5 CARBO-02 + 0.5 METHA-01 + 0.56 ETHYL-03 + 0.5 CARBO-01 + 0.5 HYDRO-01 + FORMA-01 + METHA-02 + 0.5 ETHAN-01 + 0.05 WATER + 0.2 CARBO-03 LIGNIN-S > 0.185 C11H1-01 + 0.47 P-CUM-01 + 0.5 PHENO-01 + 0.5 ETHYL-03 + 0.01 WATER + 0.1 METHA-01 + 0.1 CARBO-02 + 0.1 CARBO-01 + 0.1 HYDRO-01 + 0.1 CARBO-03 LIGNIN-H > 0.6129 ETHYL-01 + ACETO-01 LIGNIN-G > 0.9754 PROPY-01 + 0.1 CARBO-02 C11H1-01 > 0.37 P-CUM-01 + 0.601 PHENO-01 + 0.3748 PYRUV-01 + 0.1 WATER + 0.6 METHA-01 + 0.6 ETHYL-03 + 0.5 CARBO-01 + 0.5 HYDRO-01 + 0.4 CARBO-03 C11H1-01 > 0.63 C11H1-03 + 0.1 WATER + 0.5 METHA-02 + 0.45 METHA-01 + 0.2 ETHYL-03 + 1.58 CARBO-01 + 0.69 HYDRO-01 + 0.15 CARBO-03 C11H1-03 > 2.2124 PHENO-01 C11H1-03 > 0.1 WATER + CARBO-01 + 0.2 FORMA-01 + 0.7 METHA-02 + 0.7 ACETA-01 + 0.51 ACETO-01 + 0.6 METHA-01 + 0.65 ETHYL-03 + 0.5 HYDRO-01 + 5 CARBO-03

n 13

8 × 10 8 × 107 4 1 × 109

3 3 × 109 1 × 1010

1

1

Ea (kcal/mol) 52.9 36.8 10.0 30.0 11.0 2.7 3.3

4x1015

48.5

2 × 1013 1 × 109 5x108

37.5 25.5 31.5

3 × 108

30.0

7 × 103 1.2 × 109

12.0 30.0

the pyrolysis first stage and syngas for power generation. Studies on gasification process using Aspen Plus have been reported. The process is generally represented in two to four stages, particular for biomass feedstock. Mitta, et al. [26] reported simulation of a tyre gasification for syngas production with drying, devolatilisation and gasification as the main process stages, which were represented respectively by the RSTOIC, RYield and RGibbs module in the Aspen Plus. The details of the devolatilisation product from the RYield was not provided. This is very key in predicting the gas composition from the RGibbs reactor. Pala, et al. [27] reported gasification of biomass in Aspen Plus model with RYield and RGibbs as the main gasification process in the flowsheet. The RYield was used to implement biomass decomposition into elemental composition, steam and ash without accounting for solid carbon residue. Report by Lan, et al. [28] on biomass gasification in Aspen Plus shows RYield and RGibbs as the major unit process of the gasification. RYield was used to model biomass disintegration into oxygen, hydrogen, sulphur, carbon, nitrogen and ash. It was not clear whether or not a bone dry biomass feedstock was assumed. It therefore necessary to state the full composition of the feedstock used in the simulation for accurate representation of product yield from the RYield reactor. Studies have been reported that wet biomass material with moisture content up to 50% promotes syngas production with high yield of hydrogen via gasification technology [29,30]. Elimination of drying step in the gasification process could reduce the energy requirement and eventually lower the cost of production. In this study, gasification section is represented by five (5) Aspen model units: RYield (SYRDCM), RGibbs (SYNR), cyclone (SYCYC), MHeatX (GHXC) and tar separator (TARSEP). The SYRDCM reactor disintegrates the SYFD into carbon (C), hydrogen (H2), nitrogen (N2), sulphur (S), oxygen (O2), ash and steam (H2O) by specifying the yield (on weight basis) according to the proxanal, ultanal and sulfanal attributes of the feedstock (Table 1). These components are further subjected to gasification at 600–1200 °C and 0.034–1.350 air/fuel (w/w) in the SYNR reactor and syngas (SYNPD)solid particle (ash) separation is achieved in the SYCYC. The particle free syngas is quenched in the gas exchanger GHXC and tar from the gas

mechanisms of this process has been reported in the literature. According to Mattos, et al. [22], the major reaction pathway for hydrogen production from ethanol reforming is according to reaction (1) in Table 4, which produces a high hydrogen yield in six-fold from a mole of ethanol. This reaction is said to be endothermic in nature and thus temperature dependent [23] and the kinetic is first order with activation energy of 15.48 kcal/mol. [24]. The process condition used in the UPRXN were carefully chosen at 200–250 °C and 2 MPa to reduce the energy requirement for the ethanol reforming reaction and favour liquid-phase hydrogenation of furfural to furfurylalcohol [25]. Subsequently, the product stream is passed through a cooler (UPCD) to a separator (SEP-1) where unconverted sugars and reaction water product (ACQ-1) is separated from the other organic components (OG-1). In the second stage of the hydroprocessing, OG-1 stream, mainly ligninderived components and the organic product from the stage, is combined with second stream of ethanol (ETOL-2) in UPRXN-1 reactor. Shemfe, et al. [16] reported second stage hydrodeoxygenation process with pure hydrogen gas in Aspen plus using R-Yield reactor by specifying pseudo-components such as light non-volatile, heavy non-volatile, phenolics, aromatics and alkanes, Coke, H2O and outlet gases. In this study, the process is implemented using R-Gibbs reactor at 350 °C and 5 MPa since most of the reaction kinetics and stoichiometry are currently still sketchy for the pyrolytic oil upgrading, particularly for the in-situ hydrodeoxygenation reaction with ethanol as the hydrogen donor. The R-Gibbs reactor uses Gibbs free energy minimisation by calculating phase and chemical equilibria, thus requires no information on reaction stoichiometry and kinetics. Rigorous calculation method was selected with identification of possible components in the product stream based on the relevant literature. The product stream is condensed in a condenser (UPCND-2), followed by component separation in a separator (SEP-2). 2.1.3. Gasification section In the gasification section, 30% of the Napier grass bagasse (NPGBGS) that is, syngas feed (SYFD) was used to provide the heat for Table 4 Upgrading reaction 1 (Mahajani, 2003; Mattos et al., 2012; Patel et al., 2013). S/No

Reaction

k (1/s)

Ea (kcal/mol)

1 2 3

ETHAN-01 + 3 WATER > 6 HYDRO-01 + 2 CARBO-02 D-XYL-01 > FURFU-01 + 3 WATER FURFU-01 + HYDRO-01 > FURFU-02

2.50 × 104 1.79 × 106 2.0 × 104

15.48 16.36 9.24

5

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stream is recovered in the TARSEP. The final syngas product (CSYN1) is used as fuel in the combustion section.

2.1.4. Combustion section Solid fuel combustion process consists of fuel drying, decomposition, combustion and flue gas cleaning units. This is implemented in Aspen Plus in most cases with the aid RYield as decomposition reactor, RGibbs reactor and cyclone for fuel combustion and flue gas cleaning respectively [31]. The combustion section in this study used RYield reactor (CHDCMP) for char (from the pyrolysis process) decomposition into C, H2, N2, S, O2 and ash. This stream is combined with the pyrolysis non-condensable gas (NCG) and syngas (CSYN1) from the gasification section in a fuel mixer (FMX) and combusted in the RGibbs reactor (CMBRXT) with excess (AIR3) for complete combustion. Co-processing of fuel or fuel intermediates have been reported to have strong synergistic effect, which in turn increase the reactivity during conversion [32]. Heat of decomposition (QCMB) from the CHCMP reactor is used as heat source to produce high temperature flue gas (FGS). Cyclone (CMCYC) separates the ash/dust particles from the FGS and the ash-free flue gas (FGS-1) at 1286 °C is then used for steam production in the power generation section.

Fig. 3. Methodology for estimation of capital investment.

3. Results and discussion 2.1.5. Power generation section Biomass processing via thermochemical conversion route with integrated power generation system provides additional means of reducing electrical and thermal demands in the process. This approach is sometimes referred to as simultaneous thermal and mechanical/electrical energy production from a single type of fuel through combustion and subsequent energy recovery from the exhaust gas [33–35]. Herein, power generation using the flue gas from the combustion of the mixed fuel in the previous section is accomplished through steam generation from the hot flue gas with the aid of heat exchanger (STHX), steam turbine (TURB), condenser (COND) and pump (PMP). The hot flue gas exchanged heat with water (CWT) in STHX and generated superheated steam at 693 °C, 15 MPa. While the flue gas (CFG) exits at temperature of 65 °C, the high pressure steam (HPSTM) is charged into turbine (TURB) implemented at 90% and 95% isentropic and mechanical efficiency to generate electricity (POWER). The low pressure steam (LPSTM) from the TURB is then passed through a condenser (COND) to produce water (CWT1) at 53 °C.

3.1. Pyrolysis model validation Fast pyrolysis result from the simulation study is presented in Fig. 4. Effect of temperature on the pyrolysis product distribution was evaluated at 400–500 °C. Pyrolytic oil increased with increasing temperature from 47 to 68 wt% while bio-char decreased from 36 to less than 10 wt%. The non-condensable gas yield was around 15 wt% throughout the investigated temperature range. The result showed that a maximum pyrolytic oil of 68 wt% can be obtained from fast pyrolysis Napier grass bagasse at around 480 °C. Simulation result from biomass fast pyrolysis in a fluidised bed system reported by Peters, et al. [18] showed oil yield of 66.28, 69.78 and 65.98 wt% at 470, 520 and 570 °C respectively, which is in good agreement with the simulation result obtained in this study. Similarly, the result is also comparable to pyrolytic oil yield from a typical laboratory fast pyrolysis system. According to Bridgwater [36], pyrolytic oil, bio-char and non-condensable gas of about 80, 12 and 13 wt% is produced from biomass fast pyrolysis at temperature of about 500 °C under high heating rate and short vapour residence time of approximately 1 s. Recent study on fast pyrolysis of Napier grass in circulating fluidized bed reactor by Suntivarakorn, et al. [10] revealed that a temperature of 480 °C is required for maximum pyrolytic oil production. This shows that the pyrolytic reaction models used in this study to greater extent are in consonant with experimental data in the literature. Chemical composition of the pyrolytic oil from the simulation include phenolics, sugars, aldehydes, ketones, hydrocarbons and reaction water (Table 6). These components decreased with increasing pyrolytic temperature except for the reaction water, which shows continuous increase from 8.57% at 400 °C to 30.57 and 31.99% at 480 and 500 °C. This suggests a strong dehydration reaction. Similarly, the decrease in other group of organic components is evidence of cracking reactions with increasing temperature. This is also true for sugars and reaction water in the study reported by Peters, et al. [18]. The authors recorded 28.64% of pyrolytic oil as water at 470 °C, which increased to 30.58% at 550 °C with corresponding sugar yield of 30.46 and 7.68%. On the other hand, there was no clear pattern of impact of temperature on other organic compound reported. Comparing the result of simulation with experimental investigation reported in literature, high pyrolysis temperature promotes production of light oxygenates through dehydration ring-opening and fragmentation reactions [37]. Consequently, the model adequately described the distribution of pyrolytic oil composition. The summary of the mass balance across each unit in the

2.2. Energy analysis, economic evaluation and sensitivity analysis Energy analysis of the process design was implemented by assigning appropriate utility streams with the aid of Aspen Energy Analyser (AEA) tool. The corresponding energy target values are computed according to the utility load allocation by ensuring the minimum utilities needed meet up the process stream requirement. A model-based estimation approach was used to generate capital cost estimate (Fig. 3), operating cost estimate and other investment analysis using Aspen Process Economic Analyser (APEA) according to the first quarter 2016 pricing basis. The investment analysis input parameters used (Table 5) were carefully chosen to evaluate possibility of implementing the project at grassroots level particularly in developing countries. Net Present Value (NPV), discounted pay-out period (DPP), profitability index (PI) and internal rate of return (IRR) computed from revenues, capital and operating costs were used as indicators for the economic. Similarly, sensitivity analysis was performed to evaluate the impact of economic parameters on the process using Monte Carlo simulation. Effect of purchased equipment cost, instrumentation, operating labour, utilities, maintenance and feedstock on the NPV were evaluated by varying each of these parameters within ± 20%.

6

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Table 5 Investment analysis input parameters. Name

Units

Period Description Number of Weeks per Period Number of Periods for Analysis Tax Interest Rate/Desired Rate of Return Economic life of Project Salvage Value (Percent of Initial Capital Cost) Depreciation Method Escalation Parameters Project Capital Escalation Products Escalation Raw Material Escalation Operating and Maintenance Labour Escalation Utilities Escalation Project Capital Parameters Working Capital Percentage Operating Costs Parameters Operating Charges Plant Overhead General and Administrative Expenses Facility Operation Parameters Facility Type Operating Mode Operating Hours per Period Process fluid

Weeks/period Percent/period Percent/period Period Percent

1.5 2.5 1.5 1.5 1.5

Percent/period

5.0

Percent/period Percent/period Percent/period

1.0 1.0 1.0

Hours/period

Reactor Temperature (oC) 480 28.8467 37.2280 0.0099 1.6722 1.6670 30.5763

Biomass Processing Facility Continuous Processing-24Hours 8000 Liquids, gases and solids

reforming of ethanol into hydrogen with CO2 as major co-product. Nearly 30% conversion of ethanol was achieved with corresponding increase in hydrogen production from 0.13 kmol at 200 °C to 1.5 kmol at 250 °C. low temperature catalytic reforming of ethanol over platinum catalyst reported by Dai, et al. [38] showed that the major reaction in catalytic reforming is the hydrogen production pathway. They recorded about 68.1% ethanol conversion and 41% hydrogen production at initial reaction temperature (200 °C) while at 350 °C, complete conversion of ethanol was achieved. They further noted that at higher temperature, perhaps above 250 °C, decomposition of acetaldehyde is most likely. Consequently, the process variables used in this study is to allow liquidphase hydrogenation of furfural to furfurylalcohol. The aqueous phase product from the reactor mainly reaction water and unconverted sugars decreased with increasing temperature. At 200 °C, the xylose was 3.69 kg in the product stream and reached 96% conversion at the final reaction temperature (250 °C) with corresponding 94% yield of furfural. Similarly, the organic phase production increased by 40%, which is probably due to simultaneous dehydration reaction of sugars to furfural and subsequent conversion to furfurylalcohol via hydrogenation reaction. This observation is further evidenced in the linear relationship between the organic phase yield and hydrogen production. Similar trend was also noted in the liquid-phase catalytic hydrogenation of furfuraldehyde to furfuryl alcohol over Pt-catalyst reported by Vaidya and Mahajani [25]. The reaction models used in the simulation for the mild upgrading of pyrolytic oil with in-situ hydrogen production from ethanol well-described the reaction pathways. The second stage of the pyrolytic oil upgrading was modelled in Gibbs reactor since this involves some multicomponent reactions with kinetics currently not well established. The Gibbs reactor calculates the phase and chemical equilibria by minimising the Gibbs free energy. Effect of temperature and pressure on the product distribution were evaluated. At 5 MPa reactor pressure, the yield of gasoline (GSLN) product declined with increasing reactor temperature from 250 to 400 °C, which corresponds to about 18% decrease (Fig. 6). The yield of light gas product (LGP) mainly methane (CH4) and carbon monoxide (CO) with traces of hydrogen increased with temperature while the gas product stream (GP1) principally carbon dioxide (CO2) declined by 59%. These observations suggest that at reactor temperature below 350 °C decarboxylation reaction dominates the process while

Table 6 Pyrolytic oil composition from the simulation.

400 37.9780 49.0420 0.0035 2.2028 2.1956 8.5780

Year 52.0 25.0 5.0 2.0 25.0 20.0 Straight line

Percent/period Percent/period Percent/period Percent/period Percent/period

Fig. 4. Effect of temperature on pyrolysis product distribution.

Composition (wt %) Sugar Phenols Aldehydes Ketones Hydrocarbons Water

Item

500 28.2570 36.4552 0.0283 1.6390 1.6324 31.9881

pyrolysis section is presented in Table 7. 3.2. In-situ hydrodeoxygenation process Upgrading of pyrolytic oil from pyrolysis section was further upgraded in two-stage in-situ hydrodeoxygenation process. In the first stage, mild hydrotreating was conducted with hydrogen production from ethanol. Fig. 5 shows effect of temperature on the mild catalytic upgrading process and product distribution. Simulation result showed that increasing reaction temperature from 200 to 250 °C increased the 7

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Table 7 Mass flow across each unit in pyrolysis section. Material Stream Name Total Stream Mass Flows (kg/hr) WATER DEXTR-01 D-XYL-01 ETHYL-01 PROPY-01 C11H1-01 NPGBGS NITRO-01 OXYGE-01 HYDRO-01 METHA-01 CARBO-01 CARBO-02 CARBO-03 LEVOG-01 GLYOX-01 ACETA-01 ACETO-01 GLYCO-01 5-HYD-01 METHA-02 PHENO-01 P-CUM-01 FORMA-01 ETHYL-03 PYRUV-01 C11H1-03 CELLU HCELL LIG-H LIG-G LIG-S

PYFD

AIR2-2

AIR3-2

AIR2-3

AIR3-3

DPYFD

PYOIL1

BCHAR

NCG

34.4819 15.4890 – – – – – 18.9930 – – – – – – – – – – – – – – – – – – – – – – – – –

150.0000 – – – – – – – 115.0620 34.9376 – – – – – – – – – – – – – – – – – – – – – – –

200.0000 – – – – – – – 153.4170 46.5834 – – – – – – – – – – – – – – – – – – – – – – –

170.6450 18.6971 – – – – – 1.9486 115.0620 34.9373 – – – – – – – – – – – – – – – – – – – – – – –

200.0000 – – – – – – – 153.4170 46.5834 – – – – – – – – – – – – – – – – – – – – – – –

17.9741 0.9287 – – – – – 17.0444 0.0008 0.0003 – – – – – – – – – – – – – – – – – – – – – – –

14.4855 4.6337 0.0009 4.0854 0.3757 1.7741 0.6556 – – – – – – – – 0.0069 0.0016 0.0003 0.2374 0.0016 0.0007 – 0.7932 1.6820 0.0000 0.2365 – – – – – – –

3.1596 – – – – – – – – – – – – – 3.0796 – – – – – – – – – – – – – 0.0798 0.0001 – – –

2.1887 – – – – – – – 2.0000 – 0.0033 0.0260 0.0454 0.1141 – – – – – – – – – – – – – – – – – – –

Fig. 5. Effect of temperature on mild (first stage) upgrading1 at 2MPa.

decarbonylation reaction and possibly demethylation control the process at temperature above 350 °C. The composition of LPG and GP1 signifies that the second stage upgrading process is dominantly demethylation, decarboxylation, decarbonylation and hydrogenation reactions. Experimental investigation of in-situ hydrodeoxygenation of pyrolytic oil over Pd/C reported by Cheng, et al. [39] revealed that higher temperature between 250 °C and 300 °C result in decrease of hydrocarbon yield and promote gas production. The authors also reported that CH4 and CO were the main gas components at higher temperature. These observations from experimental investigation is in good agreements with the simulation results obtained. Fig. 7 shows the impact of reactor pressure on the product distribution. Increasing reactor from 1 to 5 MPa showed significant impact on the product

Fig. 6. Effect of reactor temperature (second stage) on product distribution at 5 MPa.

distribution. The yield of GSLN product increased by 10% while only 5% rise in the yield was observed between 5 and 10 MPa. Similarly, the GP1 appreciates by 53 and 14% under 1–5 MPa and 5–10 MPa reactor pressure. The LGP on the other hand, declined by 43 and 26% at 1–5 MPa and 5–10 MPa respectively. These implied that lower reactor pressure promote the yield of hydrocarbons and selectivity since increased in the production of GP1 mainly CO2 was observed, indicating decarboxylation reaction is favoured at lower pressure. Recent reviews on hydrodeoxygenation process of bio-oil suggests that lower hydrogen pressure is desired in terms of process efficiency and selectivity, which 8

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trend continued till 900 °C where the syngas attained maximum value with 50 and 45 vol% as H2 and CO value respectively. The high H2 content could be linked to the moisture content of the biomass feed. In this study, Napier grass bagasse with about 45 wt% moisture was used. According to Hu, et al. [30], in-situ generation of steam atmosphere from wet biomass feed promotes steam reforming reaction and the water gas shift reaction towards the direction of more H2 production. Methane (CH4) on the other hand declined from 7 vol% at 600 °C to nearly zero at 900 °C, which is as a result of methane steam reforming reaction that subsequently lead to more hydrogen production. Higher heating value of the syngas increased from 8.7 MJ/m3 at 600 °C to 11.5 MJ/m3 at 900 °C. Further increase in the gasification temperature showed no effect on the HHV, which is linked to the fact that the syngas attained maximum stable composition at this temperature. Mitta, et al. [26] reported increase in CO and H2 and decrease in CH4 and CO2 contents of syngas from gasification tyre when gasification temperature was varied between 750 and 1100 °C. This was attributed to probable exothermic steam methane reforming and CO2 reforming reactions. Similar observation on the impact of gasification temperature on syngas composition has been reported by Lan, et al. [28]. The authors recorded increase in CO, H2 and decrease in CO2 content of syngas as temperature was raised from 650 to 850 °C, which was validated with experimental result under similar condition. Effect of fuel/air ratio (FAR) on the gasification process at 900 °C is presented in Fig. 9. The simulation result shows that increasing FAR increased the CO, H2 content of the syngas as well as HHV particularly under FAR of 1-5. The result further indicates that FAR of 8 is needed for near maximum value H2, CO and HHV. This trend is also in consonant the observation reported by Mitta, et al. [26]. The simulation result therefore also showed that minimum FAR and gasification temperature of 8 and 900 °C is required for maximum syngas quality and HHV. The mass flow across the units in the gasification process in presented in Table 9.

Fig. 7. Effect of reactor pressure (second stage) on product distribution at 350 °C.

is also desirable for industrial scale process. It was further stated that reaction condition at 300 °C and atmospheric pressure maximizes the deoxygenation reaction [40]. Report by Wang, et al. [41] on hydrodeoxygenation of pyrolytic oil model compound also indicates that hydrogen pressure within 5 MPa is critical to deoxygenation reaction and product selectivity. These observations from the experimental investigations in literature strong agree with the simulation result. The composition of GLN product include cyclopentane, 1,3-cyclohexadiene, methylcyclopentane, traces of n-octane and cyclooctane. Summary of the material balance for the pyrolytic oil hydroprocessing is presented in Tables 8a and 8b. 3.3. Gasification process

3.4. Combustion process In the gasification process, impact of process variables such as gasification temperature and fuel/air ratio on the syngas composition and heating value were evaluated to determine the accuracy of the simulation result by comparing with experimental data in the literature. Increasing temperature from 600 °C (Fig. 8) resulted in continuous decline in CO2 while the H2 and CO increased significantly [42]. This

Effect of fuel/air ratio (FAR) on the flue gas composition from the combustion process is presented in Fig. 10. The simulation result showed that increasing FAR from 0.09 increased the CO2 content in the flue gas and became maximum at FAR of 0.21, indicating the air supplied is sufficient for complete combustion. At this point, the

Table 8a Overall mass balance of hydroprocessing section: Hydro processing section -1. Material Stream Name

PYOIL1

ETOL1

PD-C1

VP

OG-1

ACQ-1

Total Stream Mass Flows (kg/hr) WATER DEXTR-01 D-XYL-01 ETHYL-01 PROPY-01 C11H1-01 ETHAN-01 HYDRO-01 CARBO-02 LEVOG-01 GLYOX-01 ACETA-01 ACETO-01 GLYCO-01 5-HYD-01 PHENO-01 P-CUM-01 ETHYL-03 FURFU-01 PALLA-01

14.4855 4.6337 0.0009 4.0854 0.3757 1.7741 0.6556 – – – 0.0069 0.0016 0.0003 0.2374 0.0016 0.0007 0.7932 1.6820 0.2365 – –

0.1500 – – – – – – 0.1500 – – – – – – – – – – – – –

0.2806 – – – – – – – – – – – – – – – – – – – 0.2806

14.9161 5.9969 0.0009 0.1413 0.3757 1.7741 0.6556 0.1017 0.0127 0.0922 0.0069 0.0016 0.0003 0.2374 0.0016 0.0007 0.7932 1.6820 0.2365 2.5242 0.2806

8.3916 – – – 0.3757 1.7741 0.6556 0.1017 – – 0.0069 0.0016 0.0003 0.2374 0.0016 0.0007 0.7932 1.6820 0.2365 2.5242 –

6.1390 5.9969 0.0009 0.1413 – – – – – – – – – – – – – – – – –

9

GP

0.0922 – – – – – – – – 0.0922 – – – – – – – – – – –

H-1

CATR

0.0127 – – – – – – – 0.0127 – – – – – – – – – – – –

0.2806 – – – – – – – – – – – – – – – – – – – 0.2806

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Table 8b Overall mass balance of hydroprocessing section: Hydro processing section -2. Material Stream Name

OG-1

ETOL2

PD-C2

H-1

VP-1

GP1

CATR-1

LGP

GSLN

Total Stream Mass Flows (kg/hr) ETHYL-01 PROPY-01 C11H1-01 ETHAN-01 HYDRO-01 METHA-01 CARBO-01 CARBO-02 LEVOG-01 GLYOX-01 ACETA-01 ACETO-01 GLYCO-01 5-HYD-01 PHENO-01 P-CUM-01 ETHYL-03 CYCLO-01 METHY-01 FURFU-01 1:3 C-01 N-OCT-01 CYCLO-04 CYCLO-05 PALLA-01

8.3916 0.3757 1.7741 0.6556 0.1017 – – – – 0.0069 0.0016 0.0003 0.2374 0.0016 0.0007 0.7932 1.6820 0.2365 – – 2.5242 – – – – –

1.0500 – – – 1.0500 – – – – – – – – – – – – – – – – – – – – –

0.1794 – – – – – – – – – – – – – – – – – – – – – – – – 0.1794

0.0127 – – – – 0.0127 – – – – – – – – – – – – – – – – – – – –

9.6337 – – – – – 0.5873 1.4714 2.0423 – – – – – – – – – 0.0485 5.1242 – 0.1806 – – – 0.1794

2.0423 – – – – – – – 2.0423 – – – – – – – – – – – – – – – – –

0.1794 – – – – – – – – – – – – – – – – – – – – – – – – 0.1794

2.0587 – – – – – 0.5873 1.4714 – – – – – – – – – – – – – – – – – –

5.3534 – – – – – – – – – – – – – – – – – 0.0485 5.1242 – 0.1806 – – – –

3.5. Power generation

combustion efficiency recorded was 99.93%. Trace of nitric oxide (NO) as the only NOx was also detected in the flue gas under the complete combustion region due to the high temperature (excess 1286 °C) and nitrogen from air [31] while SOx (SO2, SO3) was nearly zero throughout the combustion process. When FAR was increased from 0.21 to 0.36, CO was detected in the flue gas stream with CO2 declining to 12 vol%. The corresponding combustion efficiency dropped to 58%, suggesting incomplete combustion probably due to dilution effect of nitrogen present in the air [43]. The material flow across this section of the process is summarised in Table 10.

The hot flue gas generated superheated steam at 693 °C, 15 MPa, which was charged into turbine (TURB) at 90% and 95% isentropic and mechanical efficiency to generate electricity (POWER). The simulation result indicate about 22 kW of electric power is produced from 42 kg/hr of steam. This shows that the additional electricity generation presents incentive for energy efficiency of the overall process particularly for large scale biomass processing facility. 3.6. Energy analysis The integrated biofuel production developed herein via pyrolysis

Fig. 8. Effect of gasification temperature on syngas composition and HHV at 1 kg/hr air (14.78 fuel/air ratio at 20 °C). 10

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Fig. 9. Effect of fuel/air ratio on syngas composition and calorific value at 900 °C.

and instrumentation cost which is about 43% of the total capital cost. This is followed by the purchased equipment and equipment settings (30%), and then design, engineering and procurement cost (15%). The operating cost components of the plant include the operating labour, utilities, maintenance, raw material, plant overhead, and general and administrative cost as shown in Fig. 13. The labour cost represents the highest expenses with 48%, followed by utilities (19%) and maintenance (14%). This plant requires six (6) operators per shift and one (1) supervisor at the rate of $1.5/hr and $2.5/hr per head, which resulted to overall total operating cost of $0.3 M/year. Products with economic values generated from the process are gasoline, light gas product and electric power generated. The revenues are highly dominated by income from gasoline stream (88%). Fig. 14 shows variation in the net present value (NPV) for 25 year project life-time. The NPV value is negative for the first seventeen (17) years, which signifies that the process is not economical feasible during the said period. NPV of zero (NPV = 0), indicates breakeven point and above this point, the plant become economical feasible, and becomes profitable from 18 years onward. This period is also referred to as discounted pay-out period (DPP) or payback period. The corresponding profitability index computed as the ratio of present value of the cumulative cash inflows to the present value of the cumulative cash outflows was 1.99. Internal rate of return (IRR) is another economic indicator, is the after-tax interest rate at which funds can be borrowed and breakeven at

with in-situ hydrodeoxygenation using ethanol as hydrogen source is energy intensive, which requires appropriate heating and cooling utility allocations. From the flowsheet, cooler, heaters and process heat exchanger were used with appropriate utilities. The overall thermal pinch result showed that the overall utilities (heating and cooling) are efficiently utilised. From Fig. 11, the total actual utility requirements equal to the target with CO2 emission of 4.87 kgCO2/h. 3.7. Economic performance Price of the process stream (Table 11) such Napier grass bagasse, gasoline, light gas product, nitrogen and catalyst were established in line with the quotation from vendors and the overall cost estimation was conducted in accordance with the mass balance result. The utility requirements were obtained from the energy analysis and their corresponding cost estimated using the 1st quarter 2016 pricing basis. Total capital and operating cost was estimated with the aid of Aspen Energy Analyser (AEA) V10 tool based on the mass and energy balances, and the equipment design parameters generated from the simulation. Total capital cost of $2.13 M is required for the overall plant. This includes costs associated with purchased equipment and equipment setting, piping and instrumentation, design, engineering and procurement, electrical and insulation, civil and steel, and contingencies. The capital cost distribution (Fig. 12) showed that the largest contributor is piping Table 9 Mass balance of syngas processing section. Material Stream Name Total Stream Mass Flows (kg/hr) WATER NPGBGS NITRO-01 OXYGE-01 SULFU-02 SULFU-03 NITRI-01 HYDRO-01 METHA-01 CARBO-01 CARBO-02 CARBO-03 ASH

SYFD

AIR

N2

AIR2-1

SYN

SOLID

TAR

CSYN1

14.7780 6.6381 8.1398 – – – – – – – – – – –

1.0000 – – 0.7671 0.2329 – – – – – – – – –

2.0000 – – 2.0000 – – – – – – – – – –

150.0000 – – 115.0620 34.9376 – – – – – – – – –

12.9305 0.0206 – 1.0170 – 0.1027 – – 0.9902 0.0149 10.8960 0.0217 – –

2.8475 – – – – – – – – – – – 2.8474 0.0001

0.1124 0.0206 – – – 0.0917 – – – – – – – –

12.8181 – – 0.9902 – – – – 0.8954 0.0149 10.8960 0.0217 – –

11

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Fig. 10. Effect of fuel/air ratio on flue gas composition in combustion process. Table 10 Mass balance of combustion processing section.

Table 11 Price of process stream.

Material Stream Name Total Stream Mass Flows (kg/ h) WATER SULFU-01 NITRO-01 OXYGE-01 HYDRO-01 METHA-01 CARBO-01 CARBO-02 CARBO-03

CSYN1

CHRDCP

NCG

MFUEL

AIR3

FGS

12.8181

3.1596

2.1887

18.1665

100.0000

119.7040

– – 0.9902 – 0.8954 0.0149 10.8960 0.0217 –

– 0.0007 0.0059 1.4022 0.1529 – – – 1.5979

– – 2.0000 – 0.0033 0.0260 0.0454 0.1141 –

– 0.0007 2.9961 1.4022 1.0516 0.0409 10.9414 0.1357 1.5979

– – 79.0000 21.0000 – – – – –

10.5138 – 82.0231 1.6337 0.0001 – 0.0053 25.5277 –

Item

Price

Napier grass bagasse ($/kg) Catalyst (Pd/C) ($/kg) Nitrogen gas ($/kg) Cooling water ($/kg) Ethanol ($/kg) Electricity ($/kWh) Fuel ($/MWh) Refrigerant ($/GJ) Gasoline ($/kg) Light gas product ($/kg) Potable water ($/m3)

0.0350 0.1250 0.1675 0.0001 0.7000 0.0775 26.8100 13.1100 0.7196 0.2300 0.0500

equipment cost as shown in the Tornado bar chart (Fig. 15a). Increasing instrumentation cost up to 10%, the system remains feasible and reaches breakeven point at 15%, and thereafter becomes unprofitable (Fig. 15b). On the other hand, decrease in the instrumentation cost up to 20% improves the NPV value significantly from $0.14 M base case to $0.32 M. In the case of installed equipment cost, the system remains feasible with increase in the cost up to 15% and reaches breakeven point at 20% while the reduction in the instrumentation cost by 20% increases the NPV to $0.27 M. NPV is less sensitive to operating labour, utility, feedstock and maintenance cost within the ± 20%. It is therefore evident that the plant will have economic advantage by changing the instrumentation and control systems from the fully automated option used in this simulation to semi-automation. This will also have direct impact on the installed equipment cost and significant reduction in the overall capital cost. At NPV = 0, the minimum gasoline selling price (MESP) of $5.81/gallon ($1.45/L) gasoline equivalent (GGE). This result is similar to the MFSP of liquid biofuel derived from corn stover biomass via pyrolysis reported by Anex, et al. [14]. The author reported MFSP of $2.00–5.50/GGE with pyrolysis route producing the low MFSP. Product value of $7.91/GGE has also been reported from a techno-economic performance analysis of biofuel production with electric power generation from biomass fast pyrolysis and bio-oil hydroprocessing. The high MFSP was attributed to contribution from additional cost of in-house power generation equipment, which only add 2.1% improvement to the energy efficiency.

Fig. 11. Energy balance: cooling and heating utilities.

the end of project life. IRR value of 8% was recorded. In order to further evaluate impacts of the major cost contributors to the capital and operating costs, sensitivity analysis of ± 20% changes in the installed equipment, instrumentation, operating labour, utilities, maintenance and the raw material costs were investigated. The instrumentation cost has the most significant impact on the NPV followed by installed 12

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Fig. 12. Capital cost distribution.

Fig. 13. operating cost components of the plant.

Fig. 14. Cumulative Net Present Value of different scenario for a 25 years project life-time.

4. Conclusion

capital cost of $2.13 M and operating cost/year of $0.3 M is required for 25-year economic life of the project with net present value (NPV) of $15.18 M at 8% internal rate of return (IRR). This standalone biomass refining plant is expected to produce fuel at minimum fuel selling price of $5.81/gallon ($1.45/L) gasoline equivalent (GGE). The sensitivity analysis revealed that the relatively high product value is driven by capital cost of the plant mainly, instrumentation. Consequently, the process will have economic advantage by changing the instrumentation and control systems from fully automated used in the model to semiautomation.

An integrated biomass refining for pyrolytic oil production and subsequent upgrading into biofuel via in-situ hydrodeoxygenation process has been developed using Aspen Plus® software and Napier grass bagasse as feedstock with capacity of 49 kg/h. The process plant economics such as capital and operating costs estimates, and other investment analysis were generated using Aspen Process Economic Analyser following first quarter 2016 pricing basis. Pyrolytic oil yield equivalent to 68 wt% of the initial feedstock could be produced at 480 °C with corresponding liquid biofuel yield of 5.35 kg/h. A total 13

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Fig. 15. Monte Carlo sensitivity analysis (a: Tornado chart; b: Spider chart).

Acknowledgments

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