Integral energy valorization of municipal solid waste

Energy Conversion and Management 180 (2019) 1167–1184

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Energy Conversion and Management journal homepage: www.elsevier.com/locate/enconman

Integral energy valorization of municipal solid waste reject fraction to biofuels

T

F.J. Gutiérrez Ortiza, , A. Kruseb, F. Ramosa, P. Olleroa ⁎

a

University of Seville, Escuela Técnica Superior de Ingeniería, Department of Chemical and Environmental Engineering, Camino de los Descubrimientos s/n, 41092 Sevilla, Spain b University of Hohenheim, Institute of Agricultural Engineering, Department of Conversion Technologies and of Biobased Products, Garbenstrasse 9, 70599 Stuttgart, Germany

ARTICLE INFO

ABSTRACT

Keywords: Supercritical water Biofuel Bio-oil Municipal solid waste Reforming Hydrodeoxygenation

Nowadays, the waste generation increases more and more, especially of the municipal solid waste. The municipal solid waste reject fraction may be valorized from an energy point of view. With this aim, a conceptual design of a process was developed by considering material and energy integration, which mainly consists of three sub-processes: fast pyrolysis of municipal solid waste to produce bio-oil, supercritical water reforming of the bio-oil aqueous phase to produce hydrogen to be used in the third section, which is the upgrading of the organic phase of the bio-oil by hydrodeoxygenation. The overall system was simulated using Aspen Plus software to achieve the highest process performance and the lowest utilities requirement. The former was assessed by the biofuel production (liquefied fuel gas, gasoline and diesel) and the net electrical power. In addition, the potential economic profitability of the plant was performed by specifying the main process units. Thus, for a feeding of 50 t/h of municipal solid waste reject fraction, a generation of a net electric power equal to 10.65 MWe and a production of 5.2 t/h biofuels (21.1% of the carbon present in the municipal solid waste) may be achieved, thus obtaining a very low gate fee (16.7 €/t) using the same industrial selling prices that those of fossil fuels and electricity in a full plant. Therefore, the process seems to be technically and economically feasible.

1. Introduction The technological advance along with the standard of living involves an increasingly high consumption of the energy resources and a higher generation of waste. In this context, new sources of energy must be found to satisfy the energy demand of society. These sources must be sustainable, i.e., renewable and friendly to the environment. Similarly, the amount of waste should be reduced and valorized, respectively. The fraction of the municipal solid waste (MSW) that usually goes to landfill or incineration, after separation of the fraction that may be reused and recycled, may be a candidate to be energetically valorized in a better way than incineration, and reducing the environmental problems due to soil contamination (landfill) and emissions to the atmosphere (incineration). MSW originates mostly from the households. The composition of MSW may vary from one country to another, and from one borough to another, depending on the collection system used by the administration. In the European Union, more than 200 million tons of MSW were generated in 2014 [1]. In the countries of the north and center of

⁎

Europe, such as Germany, the incineration and the recycling is mostly preferred, whereas the landfill is the most common option in the countries of the south of Europe, such as Spain. As two alternatives to the direct deposition of MSW in landfills, mixed collected waste may be processed in biological-mechanical treatment plants or thermal plants (usually, incineration). In the former, valuable elements are recovered and SRF/RDF may be produced from the reject fraction obtained; in the latter, heat and electricity may be produced [2]. This work focuses on the valorization of the waste reject fraction coming from a biological mechanical treatment (BMT), once the recyclable elements have been extracted, because this matter cannot be recovered and goes to incineration or landfill. Thus, around 40 wt% MSW entering a BTM installation can be recovered, whilst 60 wt% cannot be separated or reusable despite its high energy content [3]. Among the types of MSW reject fraction from BMT plants, some of them present a higher content in paper and cardboard (35%) as well as in plastics (35%), but also a high moisture content, while others have a high content in organic matter (43%), but less in plastics (12%) and

Corresponding author. E-mail address: [email protected] (F.J. Gutiérrez Ortiz).

https://doi.org/10.1016/j.enconman.2018.10.085 Received 26 July 2018; Received in revised form 23 October 2018; Accepted 26 October 2018 0196-8904/ © 2018 Elsevier Ltd. All rights reserved.


F.J. Gutiérrez Ortiz et al.

paper and cardboard (15%) with a relatively low moisture content [4]. Thus, the fractions subjected to pyrolysis mainly include paper, cloth, plastics and yard wastes [5] and food waste [6]. This study is aimed to achieving an energy valorization for the reject fraction of MSW coming from BMT plants as an alternative to deposition in landfills or incineration. The objective is the production of suitable liquid fuels for the automotive industry along with the generation of electricity. The use of these biofuels would reduce emissions of greenhouse gases, such as CO2. The process is designed to be energy self-sufficient, i.e., without needing an external fuel or electricity at steady conditions, so a high material and energy integration is required. The process can be divided into three sections: (1) fast pyrolysis (FP) of the MSW reject fraction to obtain a liquid product, which is separated into an aqueous phase and an organic phase (bio-oil); (2) bio-oil upgrading by hydrodeoxygenation (HDO), and (3) supercritical water reforming (SCWR) of the bio-oil aqueous phase and steam reforming of the off-gas to produce the hydrogen that is required in the HDO. The effect of main operating variables on the process performance is studied and a techno-economic analysis is carried out to assess the viability of the proposal. As far as the authors know, there is no other study that has been carried out on the proposed way to valorise the MSW reject fraction.

Bio-oil is easy to store and transport and can be used as a raw material in both the chemical industry and the fuels production. However, the bio-oil from fast pyrolysis has some characteristics that make it unsuitable for direct use as a fuel, such as its corrosive nature, high viscosity, limited volatility and chemical instability. The reason is the high content of oxygenated compounds, sometimes greater than 50% [14], so oxygen must be eliminated to improve its properties. There are several techniques on how to improve the properties of bio-oil, which can be physical processes (filtration or extraction with solvent) and chemical processes, with or without catalyst (reforming, gasification, cracking) [15]. The hydrodeoxygenation process (HDO) makes it possible to increase the H/C ratio and remove the oxygen, thus reducing the O/C ratio, simultaneously. Reaction (R1) represents this process in a simple way:

Cn Hm Ok +kH2

(R1)

Cn Hm +kH2 O

HDO is an exothermic process carried out at a high pressure (70–200 bar) and a mild temperature (300–400 °C) [16], depending on the feeding characteristics. A higher temperature favors the kinetics of the reactions but reduces the equilibrium conversion; the maximum operating temperature depends on the feed. A high hydrogen pressure favors the HDO reactions, because a high excess of H2 increases its dissolution into the bio-oil and, hence, the reaction rate. In addition, the H2 surplus decreases the formation of coke, thus prolonging the catalyst life [16]. Most of the research about HDO has been focused on the effect of different catalysts on the process. The active elements of these catalysts are mainly metals (such as cobalt, nickel, palladium, ruthenium, rhodium or platinum) either alone or combined with sulfur, phosphorus, carbon or nitrogen. One of the main drawbacks of the HDO process is the high amount of H2 required at a high pressure. Therefore, a source of hydrogen is required to achieve an integrated process. The aqueous phase of the pyrolysis liquid is a mixture of oxygenated compounds such as acids, alcohols, ketones and aldehydes, as above mentioned. These compounds can be used to produce hydrogen by a reforming process. For a generic compound, the overall reaction taking place in a reforming can be represented by reaction (R2):

2. Background Pyrolysis consists of a thermal degradation of organic material in the absence of oxygen in such a way that the large organic molecules are broken into smaller ones. The amount of the pyrolysis solid, liquid or gas products depends on the feed composition, heating rate, residence time and maximum operating temperature, as well as the type of reactor. In fast pyrolysis (FP), the production of liquid product is maximized with respect to other types of pyrolysis. The feed is heated very fast until reaching a temperature between 450 and 600 °C, and the gas residence time is about 2 s. Different reactor types are considered to realize the high heating rate. Some of them use the direct contact with a hot surface or with sand for an efficient heat transfer. Here the focus is on a process favored by small particles (around 2 mm) in fluidized bed systems [7]. The product of the fast pyrolysis of biomass is typically composed of 75 wt% liquid, 12 wt% solid and 13 wt% gas [8], although there are important differences in the distribution of products depending on the feeding source [9]. The liquid product from fast pyrolysis (bio-oil) consists of hundreds of different aliphatic and aromatic compounds with a significant presence of ethane, ethylene, propylene, pentane and acetaldehyde, which have an industrial value [10]. The bio-oil has completely different properties to those found in conventional crude oil, such as a high water content, acidity and lower heating value apart from being unstable at high temperatures and in long-time storage [9]. In addition, it is highly polar due to its high oxygen fraction (35–40 wt%, dry basis), unlike conventional crude oils whose oxygen concentration is few ppm [11]. Water is present in bio-oil from fast pyrolysis; how much depends on the nature of the feedstock [12]. If the water content is higher than about 30–35 wt%, two liquid phases are formed. In this way, the liquid fraction from fast pyrolysis consists of two phases, generally, by adding water: an aqueous phase that has polar organic molecules with low molecular weight, mainly acetic acid, but some other acids, aldehydes and ketones as well, and an organic phase (bio-oil) composed of nonpolar compounds, mostly groups of aromatic rings, but also aldehydes and ketones, with higher molecular weight [13]. Typically, the organic phase represents about 30–40 wt% bio-oil, with a water content between 20 and 35 wt%, whereas the aqueous phase has about 80 wt% water [12]. Both phases can be treated separately to produce different valuable products; the organic phase can be upgraded by hydrotreating to obtain fuels, and the aqueous phase contains a large amount of water and several organic compounds such as acetic acid, hydroxyacetone and phenol.

Cn Hm Ok +(n-k)H2 O

nCO+ n+

m -k H2 2

(R2)

A water-gas shift (WGS) reaction may follow the reaction (R2), where CO reacts with H2O to produce H2 and CO2. In general, the heat required for the endothermic reaction (R2) comes from the combustion of an external fuel. This way, the reactants would flow through the tubes located inside a furnace. Reforming is usually performed with steam. However, the huge energy required to vaporize the high content of water present in the bio-oil aqueous phase can be minimized by carrying out the reforming under supercritical conditions [17], where special properties of supercritical water can be applied [18]. This way, supercritical water reforming (SCWR) or supercritical water gasification (SCWG) is a thermo-chemical process that can convert wet biomass and organic wastes into carbon dioxide, carbon monoxide, methane and hydrogen without vaporizing water, thus reducing the energy requirement with respect to other reforming processes. Recently, SCWR of several model compounds, representative of the aqueous fraction of the pyrolysis liquid (acetic acid, hydroxyacetone, 1-butanol and glucose) has been studied, and the experimental results were compared with those obtained by modelling and simulation in Aspen Plus and Matlab by testing the compounds both separately [19] and mixed [20]. The three above-mentioned processes can be integrated in an overall process to valorize the reject fraction of MSW, as further explained below. 3. Methods Fig. 1 shows a conceptual block diagram of the proposed plant. The 1168



Fig. 1. Conceptual design of the process.

MSW reject fraction (MSW from now on) usually has a high moisture content, so it must be dried. Then, fast pyrolysis (FP) is carried out. A stream with the solid fraction, composed of ash and char, leaves the FP reactor. This stream has a high energy value and may be burnt to obtain part of the required energy. From the top of the FP reactor, the gas stream is cooled down quickly by a quench, thus preventing vapors from further reacting and condensing part of the stream. The incondensable gases are used as a fuel. A fraction of this condensate is used for the quench. The rest of the liquid stream is separated into an aqueous phase, with a high-water content, and an oil phase with a lower water content (bio-oil). The aqueous phase is reformed under supercritical conditions (240 bar and 800 °C) to obtain H2, although other gases are also produced, such as CO2, CO and CH4. The gas product at high pressure is expanded to generate electricity. Subsequently, two stages of water gas shift (WGS) are carried out at high and medium temperature, respectively, to increase the hydrogen production, which is separated from the other gases by pressure swing adsorption (PSA). The off-gas is partially reformed with steam, thus increasing the H2 production, which is used in the HDO of bio-oil to reduce its high oxygen content and increase the H/C ratio. The upgraded bio-oil in the HDO is distilled, obtaining different fuel streams (LPG, gasoline and diesel). Finally, the surplus hydrogen is fed to a PEM fuel cell to produce extra electricity. This is the basis of a self-sufficient biorefinery process; it needs no infrastructure but several unit operations, as well as a high material and energy integration, as shown below.

performing the mechanical-biological treatment of MSW of four provinces of western Andalusia (Spain): Seville, Huelva, Cádiz and Córdoba [22]. The value is 46.43 t/h, which has been rounded up to 50 t/h in this study. 3.1.1. Fast pyrolysis section The fast pyrolysis section (Fig. 3) includes MSW drying, pyrolysis, separation of solids, char combustion, quench of the gases and vapors, and separation of the pyrolysis liquid in two phases. First, once the feed is grinded and its size is reduced to 2–6 mm, the moisture content must be lowered to 7–10 wt% [23]. The drying of the feed was simulated with the dryer model, specifying 8 wt% of moisture at the dryer outlet and using a stream of inert gases at 300 °C and atmospheric pressure. This gas stream is the flue gas with 6 vol% oxygen, after exchanging heat in other parts of the plant. Then, the dried waste is fed into a circulating fluid bed reactor, where a fast pyrolysis is carried. A fraction of the incondensable gases without O2 is used as a fluidization agent. In the reactor, the fluid residence time is about 2 s and the reaction temperature (500 °C) is reached almost instantaneously. At the reactor outlet, the vapor and the incondensable gases are separated from the formed solid particles (inert and char) by cyclones. The solids are conveyed to another fluidized bed where the char combustion takes place, releasing energy used in the pyrolysis reactor by heat transfer through continuously recirculated inert solid particles, such silica; this latter part was not simulated. The highest oxygen content bio-oil fractions contain oxygen as carboxylic acids, carbonyls, phenols, and alcohols [24]. Due to the large number of compounds and the lack of kinetic data, the pyrolysis reactions have been modelled using a R-Yield reactor, setting the reaction product concentration based on biomass composition data [23] along with the waste elemental composition (Table 1). The result is shown in Table 2. As above mentioned, paper, cloth, plastics and yard wastes are susceptible of fast pyrolysis. Because of their main components are hemicellulose and cellulose, waste paper and cardboard are important representative constituents of biomass in MSW; similarly, yard wastes, chopsticks and used furniture contribute as wood and woody mass in MSW. In terms of components, wood and woody mass contain cellulose,

3.1. Process design and simulation Fig. 2 shows the Aspen flowsheet of the overall process. Below, the different sections in which the process is divided are described using smaller flowsheets that are easy to read and where the streams and units may be tracked down. The MSW has a very heterogeneous composition and there is great variability depending on the region considered. In this case, a representative composition of the MSW is shown in Table 1 [21]. The amount of MSW fed to the plant is an average of the reject fraction after 1169


Fig. 2. Flowsheet of the overall process.


1170



Table 1 Elemental composition of the feed.

Table 2 Composition of the fast pyrolysis product.

Elemental Analysis (dry basis), wt% C H O N S Ash (Moisture; wet basis)

Phase 46.25 6.18 40.58 0.43 0.18 6.38 (19.02)

Gas

hemicelluloses and lignin [5]. Likewise, the fast pyrolysis of food waste produces acetic acid or furfural [6]; moreover, in vegetables and fruits, there are different fibers that may be composed of hemicellulose, cellulose and lignin. Therefore, the chosen composition may be considered as representative of many MSW reject fractions to be pyrolyzed, although it must be kept in mind that this an approach is taken due to the lack of data regarding the real composition of the MSW reject fraction from fast pyrolysis. The separation of the produced solids is simulated by an ideal separator (S-Split). The char combustion is simulated by means of a RGibbs reactor. However, because of this model does not work using unconventional solids, a R-Yield reactor is previously used to deal with the char in its elemental components: C, S, N, O, and H. The air for the char combustion is tuned to achieve 6 vol% oxygen in the flue gas. Likewise, the flue gas temperature is limited to 1150 °C, to prevent mechanical problems in the materials. The combustion heat is used to heat both the waste to be pyrolyzed (at 500 °C) and the

9.2

Compounds

wt%

CO2 CO CH4 H2 Ethylene Ethane

3.15 3.94 1.14 0.41 0.43 0.13

Liquid

62.19

Water Acetic acid Formic acid Propionic acid Methanol Hydroxyacetaldehyde Acetaldehyde Formaldehyde Hydroxyacetone Furfural Levoglucosan Phenol Guaiacol Ethylthioethanol 2-Pyrrolidone

Solid

28.61

Ash Char:

Fig. 3. Fast pyrolysis section. 1171

wt%

20.68 3.47 1.35 0.27 0.6 15.07 0.11 1.03 2.98 0.54 11.6 0.38 3.25 0.05 0.81 C H O N S

5.87 20.16 0.74 1.43 0.26 0.15



fluidization gas. To avoid further reactions in the bulk of vapors and gases, a quench is carried out by recirculating a part of the pyrolysis liquid stream. The recirculated fraction is computed to reach a temperature of 90 °C after mixing [25]. After, the condensate is cooled down to 50 °C using cooling water; then, the liquid and the incondensable gases are separated in a flash unit. The incondensable gas, composed of H2, CH4, C2H6, C2H4, CO2 and CO, is used as a fuel. As above mentioned, a part of these gases is recirculated to fluidize the bed of MSW in the pyrolysis reactor, which is calculated using a design specification in which a ratio of 1.6 kg of gas per kg of waste is set [26]. Water is added to the liquid to separate it into two phases, an aqueous phase and an organic phase. The water comes from the residual stream of the reforming section. A value of 65 wt% of water in the pyrolysis liquid prior to separation was assumed, and 85 wt% of water in the stream entering the supercritical reforming section was set by a design specification. The separation of phases cannot be performed using the model of a decanter or an R-Gibbs due to the lack of interaction parameters for the calculation of the liquid-liquid equilibrium between many of the present compounds. Thus, the distribution of phases given by Oasmaa et al. [9], based on the water content in the feed, was used to obtain a regression model for the estimate of the mass fraction of water in the aqueous fraction ( X H2 O-aq ) from the mass fraction of water in the fast pyrolysis stream ( X H2 O-in ), as expressed by Eq. (1).

X H2 O

aq

=

0.2875 + (3.175X H2 O

in )

(2.25X H O in2); 2

The first stage of hydrogen production is a reforming process under supercritical water conditions (SCWR). This process takes place at 240 bar and 800 °C. The reforming reaction is endothermic, so a heat input is needed to reach and keep the reaction temperature, which comes from the combustion of the incondensable pyrolysis gases in a furnace. The reactor is modeled by an R-Gibbs. The gas product stream, composed mainly of H2, CH4, CO2, CO and H2O, is expanded by a turbine to 40 bar, thus obtaining electrical power. Afterward, two stages of WGS are carried out at 400 °C and 250 °C, respectively, to increase the H2 production; the temperature increases up to 425 °C and 252 °C, respectively. The WGS stages are simulated using two R-Gibbs reactors in series. Then, the stream enters a flash unit where is cooled down to 35 °C to condense and separate the water. The resulting gas product joins the recirculated gas leaving the HDO process along with the off-gas from the steam reforming reactor and the gas coming from the distillation section to enter a PSA system, where H2 of high purity is obtained and CO2 is separated for sequestration. The PSA system consists of six sets of adsorbers that have been modeled as ideal separators. Upstream from the system, water, H2S and NH3 are sequentially removed in three beds using different sorbents (silica gel, activated carbon and activated alumina/impregnated activated carbon, respectively), in two different sets of adsorbers. Then, in the first PSA, 95% of the H2 entering the system is withdrawn. Subsequently, different beds make it possible to separate CO2 for sequestration, pure H2 to be sent to the HDO system, N2 stream is sent to the furnace of pyrolysis gas because it has enough H2, and CO plus CH4 enter the steam reforming reactor. For the desorption of the gases in the beds the pressure must be lowered, so compression stages are needed prior to the next adsorption bed. The mixture of light hydrocarbons (C2C6) is separated from the CO2 in another PSA system, where a fraction of the rich-stream of CH4 is used for the regeneration, thus avoiding the CO2 in the distillation system, whose performance get worse if CO2 is present. The stream of light hydrocarbons (C2-C6) is sent to the distillation system, as shown in the next section. H2 from the SCWR section (including the two WGS stages) is fed into the HDO reactor along with the bio-oil. The stream of gases composed of hydrocarbons from C1 to C5 and CO, which is mostly obtained in the HDO reactor, is used to increase the H2 production by steam reforming (SR), where the steam-to-carbon molar ratio is 6 [27]. The SR is an endothermic process that takes place at 25 bar and 800 °C, and the required energy comes from the flue gases of the pyrolysis char combustion. The PSRK thermodynamic method was used to model the supercritical state, whose good results have been shown by various sources

r2 = 0.9982 (1)

Eq. (1) may be applied in the following ranges: from 0.3 to 0.6 for X H2 O-in and from 0.46 to 0.81 for X H2 O-aq , matching to mass water fractions in the organic phase that vary between 0.24 and 0.30. The thermodynamic method used for the drying, pyrolysis and char combustion was the Peng-Robinson state equation, because solids and non-polar compounds are involved. For the quench and the L-V separation, the UNIQUAC method was used, due to the presence of polar compounds and a pressure lower than 10 bar. 3.1.2. Hydrogen production section The flowsheet for the simulation of this section is shown in Fig. 4. The aqueous phase from the pyrolysis liquid is pumped up to 240 bar and preheated before entering the SCWR reactor. The preheating is achieved using different process streams, which is explained in more detail in Appendix A.

Fig. 4. Hydrogen production section. 1172



Fig. 5. HDO section.

[19]. The Peng-Robinson state equation was used in those units where gaseous hydrocarbons are present.

Table 3 Composition of the HDO product [23].

3.1.3. Hydrodeoxygenation and distillation section The bio-oil upgrading is performed by hydrodeoxygenation (Fig. 5), where oxygen is removed and the H/C ratio is increased. In the HDO process, a large excess of H2 is required, with an H2/bio-oil (water-free) molar ratio of 8 (or 0.22 for a mass ratio) [23], so unreacted H2 is recycled to keep 22 wt% of bio-oil, and the H2 surplus is fed into a PEM fuel cell. The HDO takes place at 100 bar and 300 °C, which are typical values found in the literature [14]. The reactants heating is performed using process streams, as explained in Appendix A. In the HDO, three phases coexist: liquid (bio-oil), gas (H2) and solid (catalyst). Due to the lack of kinetic data of the reactions occurring in the HDO, the simulation was performed by establishing a series of products and their distribution from data obtained in the literature, adjusting the concentration of each compound from the mass balance of the feed (Table 3). The composition is taken from literature [23], where representative model compounds and their distribution in the product of HDO for each group of compounds are provided. The catalytic HDO reactor must be cooled down to keep the reaction temperature constant. The amount of cooling water is calculated considering the reaction enthalpy. The output of the reactor is cooled to 35 °C to condense most of the product. The gas stream, which consists of H2, CO2 and light hydrocarbons (C1-C5), is expanded to 25 bar and sent to the PSA system, to recover the H2 and the hydrocarbons for the distillation system, respectively. The condensate is expanded to atmospheric pressure. The liquid is a mixture of the upgraded bio-oil (biofuel) and water, which are easily separated by decanting (decanter model). Once the water is removed, the biofuel is heated and distilled, obtaining the different biofuel fractions: LPG, gasoline and diesel. A fraction of the LPG stream is gaseous and separated in a flash, thus sending it to steam reformer in order to produce H2. Distillation was firstly modelled using a 4-column system and the DSTWU model (Fig. 6), and after refined by the more rigorous RadFrac model. A reflux equal to 1.2 times the minimum value was set and the key compounds were selected according to the boiling point of the model compounds used, taking into account the distribution assumed for these between the different fuel fractions, as follows:

Group

Model compound

wt% (dry basis)

Gases

CO CO2 CH4 Ethane Propane Butane Pentane NH3 H2S

n-Paraffins

Hexane Dodecane Octadecane

6.57 3.78 2.40

i-Paraffins

3-Methyl hexane 4-methyl nonane

2.82 3.82

Cycle C5

Ethyl cyclopentane 1-methyl-1-ethyl cyclopentane

2.35 3.11

Cycle C6

Cyclohexane Butyl cyclohexane

3.01 2.27

Bicyclos

1,1-bicyclohexyl

2.09

Cycle C7+

1,3- dimethyl adamantane

2.13

Aromatics

O-xylene 1-ethenyl-4-ethyl benzene

4.89 2.07

Phenanthrenes

4-methyl phenanthrene

3.99

Pyrenes

Pyrene

3.87

Diphenyl compounds

1,2-diphenyl ethane

0.66

Indanes e indenes

Indane 1,2,3-trimethyl indene

1.06 0.80

Naphthalenes

1-n-hexyl-1,2,3,4-tetrahydro naphthalene 2,7-dimethyl naphthalene

1.55

Polynuclear Aromatics

1-phenyl naphthalene

3.30

Oxygenates

5-methyl-2-(1-ethyl methyl) phenol

3.01

Nitrogenates

2,4,6-trimethyl pyridine

0.10

Organosulfur

Dibenzothiophene

0.02

0.66 13.18 6.25 4.92 3.45 3.46 4.65 1.42 0.07

2.67

In the HDO and distillation section, the Peng-Robinson method was used, while the UNIQUAC method was applied for the L-L equilibrium.

– Compounds lighter than C5 (included) are part of the LPG. – Compounds C6 - C10 (boiling points: 68–191 °C), with somewhat of C4, C5 and C12, form gasoline. – Compounds C12 - C18 (boiling points: 203–314 °C), with somewhat of C10, form diesel.

3.1.4. Fuel cell section The surplus H2 is fed into a PEM fuel cell, where electricity is generated due to the movement of electrons produced by the dissociation of H2 in protons and electrons with the subsequent formation 1173



Fig. 6. Distillation model with four columns.

of water from the reaction between hydrogen and oxygen. When the molecules of H2 are broken, the protons pass through a membrane until they reach O2, while the electrons travel through an external circuit, generating an electric current. Air and hydrogen must contain water to keep the membrane saturated, thus prolonging its life, and to cool the fuel cell [28]. The fuel cell was modeled as an isothermal reactor (RStoic) operating at 80 °C, where the reaction (R3) takes place:

H2+0.5O2

H2 O

carbon steel and an atmospheric pressure. Thus, CBM includes indirect expenses (freight, insurances, taxes, construction overheads, contractor engineering expenses, and so on), structural supports, piping and material required for installation, among others. In addition, the bare module cost (CBM) plus contingency and fees costs (assumed to be 15% and 3% of CBM, respectively) is the total module cost (CTM). The total investment cost (TIC) or grassroots costs is the sum of CTM and the auxiliary facilities costs, which were assumed to be equal to 50% of CBM, within the possible values indicated by Turton (between 20 and 100%) [29]. The cost of the units is given in previous years; as the cost of purchasing equipment varies with time, an update must be done to the current year. This is done using the Plant Cost Index in Chemical Engineering (CEPCI) for the reference year and for the current year (Eq. (3)):

(R3)

The electrical power of the fuel cell depends on its electrochemical performance, which is 45% [28]. In the simulation, a heat stream is obtained from the reaction (R3), and 45% of the heat generated is converted into electricity. This is modeled with an F-Split in the reactor heat stream. In the steady state, 420.6 kg/h of H2 that enter the fuel cell generate 5.13 MWe.

Cost current = Cost ref

3.2. Process economics

FA FB

(3)

All the costs were updated to 2017 (CEPCI 567.5) and an exchange rate of 1 € = 1.1998 $USA (December 31, 2017) was used. Fixed operating costs were estimated by applying percentages of the total direct costs (TDC) derived from the capital costs, based on [30]. On the other hand, variable operating costs are due to consumables (cooling water, waste water treatment, catalysts for reactors and sorbents for adsorbers); external steam and electricity are not required as they are produced in-situ. In the procedure, cash flows (revenues from sales and fee gate minus expenses) before tax (30% on profit) and after-tax were computed, by choosing a linear goods depreciation of 10 years and a negligible salvage value. Other assumptions were a plant life of 20 years, a construction/start-up period of one year, a working capital of 5% TIC, making the payment of 50% of the total capital at the beginning and the rest at the end of the first year, just when the plant is started up.

The investment required for the installation of the plant was estimated from the purchase cost of the main units. The method used is based on the procedure described in [29]. The cost of the units was obtained using the CAPCOST tool and various techno-economic reports, scaling to the capacity of the units by Eq. (2):

CA = CB ·

CEPCIcurrent CEPCIref

n

(2)

where Ci is the cost, Fi is the capacity and n is a scale factor. In the CAPCOST tool, the bare module cost (CBM) is the sum of direct and indirect costs associated with equipment purchase and installation. These costs are estimated using a multiplication factor that considers the specific manufacturing materials and the operating pressure with respect to the base unit cost, which is based on the use of 1174



Table 4 Mass balance of the fast pyrolysis and hydrogen production sections. Stream

Mass flowrate (kg/h)

MSW reject fraction fed to the plant Water removed from MSW (drying) Fast pyrolysis section Feeding to the fast pyrolysis Solid product of the pyrolysis (char + ash) Pyrolysis incondensable gas to be burnt Pyrolysis liquid product Addition of water to obtain two phases Liquid aqueous phase (to SCW reforming) Liquid organic phase (bio-oil) to HDO reactor Fluidization gas for the pyrolysis Recirculated liquid for the quench Air for the combustion of char Flue gas from the combustion of char Ash Hydrogen production section Air for combustion of the pyrolysis gas Flue gas from the combustion of the pyrolysis gas Products from SCWR

H2 H2O CH4 CO2 CO N2 H2S

Hydrogen leaving the second WGS reactor Feeding to SR (H2S: H2 0.2 kg/h) H2O CO2 CO CH4 C2-C5 NH3 Gas products from SR, after cooling from 800 °C and condensing (H2S: 0.2 kg/h)

H2 H2O CO2 CO CH4 N2 NH3

Water required for SR

Temperature (°C)

Pressure (bar)

50,000.0

25.0

1.01

5,989.0

155.7

1.01

44,010.9 12,591.5

155.7 500.0

1.01 1.01

4,844.1

50.4

1.01

26,575.8 25,340.5

50.0 30.6

1.01 1.01

32,066.3

39.2

1.01

19,850.1

39.2

1.01

64,784.0 106,303.2 155,000.0 165,008.1

500.0 50.0 25.0 1150.0

1.01 1.01 1.10 1.01

2,583.4

1150.0

1.01

28,934.3

25.0 (preheated to 300 °C) 1000.0

1.10

454.6 24,907.6 454.0 4,863.9 1,367.4 16.4 1.9

800.0

240.0

552.4

251.6

240.0

1.0 10,910.5 1,467.8 2,135.2 1,229.2 349.1 24.9

790.0

25.0

625.8 20.5 5,068.2 2,074.6 324.7 19.9 0.8

35.0

11,021.8

49.0

33,778.4

Table 5 Mass balance of the HDO process and the PSA system, as well as water removed from different units. Stage/unit

Compound

Feeding to HDO

H2 H2O CO2 CO CH4 Organic compounds NH3 H2S

2,958.6 6,749.8 148.2 4.7 1.9 12,886.1

H2 H2O CO2 CO CH4 Organic compounds NH3 H2S

3.9 12,980.8 51.4 0.3 6.7 4,592.8

Liquid product leaving HDO

Temperature (°C)

Pressure (bar)

300.0

100.0

36.2

25.0

57.3 0.4

104.5 0.8

Gas product from HDO to PSA

H2 H2O CO2 CO CH4 C2 C3 C4 C5 H2S NH3

2,208.2 12.9 950.6 49.9 468.7 347.9 211.9 152.5 106.1 4.5 3.3

35.3

25.0

Feeding to PSA

H2 H2O CO2 CO CH4 C2 C3 C4 C5 H2S

3,906.8 52.9 13,145.5 2,432.5 2,153.8 345.9 208.9 148.2 101.4 6.51

35.2

25.0

H2 from PSA

To HDO To PEMFC

2,956.9 420.6

35.2

25.0

24,312.7

35.0

25.0

16,122.3

35.0

25.0

13,137.2

36.9

25.0

1.10

Water removed from the products of SCWR + WGS Water removed from the products of SR Water removed from the products of HDO

25.0

Caudal (kg/h)

4.1. Mass and energy balance. Efficiencies

25.0

Tables 4 and 5 show the mass balance of the fast pyrolysis and hydrogen production sections as well as the HDO process and the PSA system. About 12.0 wt% of the water is removed by drying. Likewise, a large fraction of the dried MSW fed into the fast pyrolysis reactor is converted into solid (25.2 wt% MSW), and only 53.2 wt% MSW is obtained as pyrolysis liquid. The aqueous phase has 28.2 wt% total carbon of pyrolysis liquid. The organic phase has the remaining 71.8 wt%. It should be noted the high flow-rates of air, required for the combustion of the pyrolysis char, and the high flow-rate of the recirculated liquid to cool and condensate the pyrolysis vapors from 500 °C to 90 °C. The H2 surplus required for the HDO process involves a large recirculation of unreacted H2, which increases the flow-rate of gas entering the PSA system. The total net H2 production is 1178 kg/h (46.9%

4. Results and discussion In this section, mass and energy balances of all the plant are carried out. This way, some process performance parameters such as the energy efficiency, biofuels and electricity specific production, and carbon efficiency are obtained. Furthermore, the effect of the main operating parameters (feed composition, water added to obtain the two bio-oil phases and operating conditions) on the process performance is studied by a sensitivity analysis. Finally, the economic analysis of the proposed process is carried out by a discounted cash flow analysis, and the minimum gate fee is obtained.

1175



The overall and thermal energy efficiencies, defined by Eqs. (4) and (5), take into account the lower heat value (LHV) of biofuels with respect to the energy content in the feed compounds, and the overall energy efficiency also considers the net electrical power (Wnet ).

Table 6 Inputs and outputs of HDO and distillation column. Product

Mass flowrate (kg/h)

Bio-oil entering the HDO reactor Water in bio-oil H2 fed to HDO Gas product from HDO to PSA Liquid product

Distillation products

Raw biofuel (upgraded bio-oil) Water LPG (liquid product) Gasoline Diesel

Carbon efficiency (%) based on upgraded bio-oil

pyrolysis liquid

MSW

Thermal energy efficiency:

Th energy

=

i

mbiofuel i LHVbiofuel j

19,850.1 6,749.8 2,958.6 4,631.1

Overall energy efficiency:

5,195.5

Ov energy

=

Wnet +

mfeed jLHVfeed i

i

(4)

j

mbiofuel i LHVbiofuel

mfeed jLHVfeed j

i

j

(5) The values of the thermal energy efficiency and overall energy efficiency were 35.7% and 41.4%, respectively. Although these figures could be considered as moderate, the added value of biofuels will increase the process efficiency from an economic point of view, as shown below.

12,980.8 648.9

7.00

5.07

2.24

2,710.7 2,270.4

29.95 29.87

21.75 21.69

9.60 9.57

4.2. Sensitivity analysis

in the SCW reforming plus WGS stages and 53.1% in the steam reforming). The results obtained in the simulation of the HDO system and the distillation of the upgraded bio-oil are shown in Table 6. Although the HDO product has a large amount of water, this is completely separated in the decanter thanks to its immiscibility with the biofuel. Regarding the mass efficiency or yield of the process relative to the biofuel production, only 28.4 wt% of the bio-oil fed to the HDO reactor is converted into final biofuel. This is due to the water content of the bio-oil and because of the oxygen removal, which leaves the reactor as water after reacting with hydrogen. Although the above yield is low, most part of carbon of bio-oil ends up in biofuels. This is also shown in Table 6. Only 21.1% of the carbon present in the MSW is part of the biofuels produced; the rest is char and CO2 coming from reactions occurring in the plant, mainly from combustion and reforming, although a large amount is separated for sequestration (11.4 t/h, i.e., 14.6 wt% C of MSW). The carbon efficiency increases up 48.5% and 66.8% taking as references the pyrolysis liquid and the upgraded bio-oil, respectively. The remaining fraction of bio-oil corresponds to the formation of gases in the HDO process, which enter the PSA system where a part of them is sent to the SR reactor. The fraction of carbon in the products with respect to the pyrolysis liquid is lower due to the organic compounds present in the aqueous phase, which is used to produce H2. Regarding the emissions of the plant, a large part of the carbon present in the MSW is emitted with the flue gases (50.7%) coming from the combustion of char and pyrolysis gases. The separated process water has organic components that have not been completely separated from the reforming products in the flash units, which also lessen the yield with respect to the carbon in the production of biofuels. Another waste stream is the solid ashes withdrawn from the pyrolysis process, once the char combustion has taken place. Table 7 shows the energy balance of the plant. The heat integration allows the energy self-sufficiency, so the thermal energy that remains in the cooling water can be used for district heating in the plant itself or surroundings. Likewise, a net electric power of 10.65 MWe is generated, once discounted the power consumed in fans, compressors and pumps. It should be noted the influence of the power required by the compression in the PSA system, as above mentioned. Compression of H2 from 25 to 100 bar also affects the electricity balance. Nonetheless, the process adds energy to the electrical grid. This way, electricity is another important output of the plant along with the biofuels (LPG, gasoline and diesel). On the other hand, the electricity consumption has not been taken into account in some parts of the plant, such as those relative to the pressure drop throughout the pipes and units of the system, as well as in other facilities such as buildings or laboratories, which are not significant.

In order to know the effect of different parameters, such as the operation conditions (P, T) and the water content of streams in different parts of the process, on the biofuels and electricity production, a sensitivity analysis has been carried out. Only one factor was changed at-atime. Regarding the effect of the operating conditions, the H2 production in the SCWR reactor increases with the temperature up to a maximum of 600 kg/h at 1170 °C. However, beyond 900 °C where the H2 production is 550 kg/h, the increase is less and less, requiring a source of heat at high temperature, which could also cause mechanical problems in the manufacturing materials. On the other hand, the WGS reactors make it possible to increase the H2 production without using such high temperatures. Likewise, the pressure does not have a significant effect on the H2 production beyond 240 bar, i.e., at supercritical conditions. Regarding the effect of the temperature and pressure in the steam reforming process. At pressures of up to 25 bar and beyond 800 °C, the increase in H2 production is less and less. This smaller increase in H2 production takes place at higher temperatures as pressure rises (850 °C at 40 bar, and 900 °C at 50 bar). Nevertheless, temperatures higher than 800 °C involve special manufacturing materials, apart from being more energy-consuming. On the other hand, a decrease in pressure from 40 bar to 15 bar leads to a significant increase in the H2 production about 20% at 800 °C. Fig. 7 shows the effect of the pressure in the steam reforming process at 800 °C, assuming equilibrium. This difference is more significant at lower temperatures, although the mass flow-rate of H2 may become too low. A pressure lower than 25 bar could be interesting, as the H2 surplus could be sent to PEM fuel cell that would compensate the increase in electrical power consumption due to the H2 compression for the HDO process. However, that lower pressure should be moderate because there could be a kinetics limitation in the steam reforming in a real process. Likewise, the electrical power is maximum at 25 bar, and barely decreases at pressure values from 15 to 35 bar; the same occurs regarding the different heat power items involved in the SR reactor. Similarly, the operating pressure in the HDO can have a significant influence on the electrical power required for the compression of the H2 and, hence, in the net generation of electricity. Thus, varying the pressure between 40 and 150 bar, the power consumption varied less than ± 1.1 MWe, taking the pressure of 100 bar as a reference. Therefore, the operating conditions are suitable for this process. Regarding the effect of the water content in the feeding to the SR reactor on the production of H2, CO and CO2, the greater the water content, the greater the H2 production, as in the SCWR process. Likewise, CO yield is almost constant while the CO2 production increases as the water content is increased. However, a more interesting parameter is the water added in the separation process of the bio-oil 1176



Table 7 Energy balance of the main parts of the plant. Heat flows entering/leaving the main units (kW)

Heat exchangers – CODE (kW)

MSW heating and reaction hold at fast pyrolysis reactor Gas fluidization heating Heating up to 800 °C and reaction hold at SCWR reactor Heating up to 800 °C and reaction hold at SR reactor Released heat from HDO (at 300 °C) Condensers (4) of distillation column

Heat exchangers – CODE (kW) Liquid pyrolysis cooling Air preheating (pyrolysis gas combustion) Gas preheating (pyrolysis gas combustion) Aqueous phase preheating for SCWR

18,705.3 15,923.2 15,327.6 4,118.2 −9.960.0 −509.5

HE01 HE02 HE03 HE04 HE05 HE06 HE07 HE08

Preheating of SR feeding Cooling of steam reforming product Preheating of HDO feeding Cooling of HDO product Cooling of feeding to second column C.W. heating to increase the steam fraction C.W. heating to increase the steam fraction Steam superheating H2 preheating for PEMFC 1st condensation of steam from turbine Heating of water leaving the plant 2nd condensation of steam from turbine Gasoline cooling to 35 °C Diesel cooling to 35 °C

34,221.7 2,212.5 1,190.3 7,017.8 3,475.8 2,171.6 4,281.5 920.1

Enthalpy of the inlet streams (kW) MSW reject fraction Air for char combustion Air for pyrolysis combustion Cooling water at 25 °C Refrigerated water at 5 °C Air for PEMFC Water for PEMFC

57.4 348.0 35.7 4,401.7 2,278.0 52.5 428.0 80.6 81.8 87.2

TOTAL

7,850.9

coming from the pyrolizer to obtain the oil phase and the aqueous phase. Based on the data reported by [9], the process cannot be energy self-sufficient if the water content in the aqueous phase was 87 wt% instead of 85 wt% (base case). By decreasing the water content in the aqueous phase from 85 to 70 wt% (the water added is reduced from 25.4 t/h to 3.3 t/h), the H2 production decreases about 22%, the electrical power production diminishes from 18.50 to 14.25 MWe, but the bio-fuels production, the distribution among LPG, gasoline and diesel barely change (< 0.5%). Finally, the effect of a change in the MSW elemental composition on the process performance was carried out (Fig. 8). If a higher ash content were present in the feed, the net electrical

HE09 HE10 HE11 HE12 HE13 HE14 HE15 HE16 HE17 HE18 HE19 HE20 HE21 HE22 HE23 HE24 HE25 HE26 HE27

4,970.6 7,934.2 8,580.5 5,563.4 572.3 6,870.7 9,534.5 294.1 4,552.4 553.4 985.0 26,638.4 5,011.2 74.2 36,867.0 332.7 10,798.6 152.5 369.0

Enthalpy of the outlet streams (kW) LPG Gasoline Diesel Waste water Flue gas emitted to atmosphere Flue gas emitted after MSW drying Gas leaving the PSA (CO2) H2S, NH3 and H2O leaving the PSA Water to district heating Ash

102,672.2 11.6 2.3 4,670,391.0 44,681.6 1,854.0 1,915.2

Electric power entering the system (kW) – consumption – Pump of bio-oil Pump of aqueous phase Rest of pumps Compressors (five) of PSA system with intermediate cooling Compressor for H2 Compressor for LPG (distillate leaving the column C-4) Fan of air for char combustion Fan of air for pyrolysis gas combustion Fan of air for PEM Fuel Cell Chiller for refrigerated water (5 °C)

533.1 1,222.6 329.1 39,681.2 29,810.0 107,480.0 28,215.0 238.3 4,620,685.9 309.2

Electric power leaving the system (kW) – generation – Turbine for product of SCWR Steam Turbine PEM Fuel Cell

−5,517.9 −7,851.6 −5,131.1

TOTAL

−18,500.6

power and biofuels production would be clearly lower. Thus, if the elemental composition (wt%) were 35.81% C, 4.82% H, 24.43% O, 0.78% N, 0.41% S and 33.75% ash, the net electrical power and the biofuel production would be 6.82 MWe (a reduction of 3.43 MWe) and 2533 kg/h (a decrease of 50%), respectively. Likewise, if the MSW feed was richer in carbon and hydrogen (55% C, 8% H), with less content in oxygen and sulfur (30% O, 0.04% S), keeping the same nitrogen and ash as in the base case (0.43% N, 6.5% ash), the net electrical power would increase up to 13.15 MWe but the biofuel production is lower 3779 kg/h (a decrease of 25%), especially gasoline and diesel. This later takes place because more char is produced (almost 15 t/h) and pyrolysis liquid production is lower (20.33 t/h) and, hence, the bio-oil and

9 000

640 630

8 000

620 610

250

600 590

580 570 560

Q SR (kW) Q preheating (kW) Q cooling (kW)

6 000 5 000

540

15

20

25

Pressure (bar)

30

35

200

W feeding compressor (kW) W feeding pump (kW)

150 100 50

4 000

550

Power (kW)

7 000

Q (kW)

H2 production (kg/h)

Product cooling for SCWR + WGS

0

3 000

15

20

25

Pressure (bar)

30

35

15

20

25

Pressure (bar)

Fig. 7. Effect of the pressure in the SR reactor on the H2 production, heat power and electrical power at 800 °C. 1177

30

35



FP production (kg /h)

20000

15000 10000

5000

35000

1400

Base case Alternative 1: higher amount of ash

30000

1200

Alternative 2: more C, H; less O, S; same N, ash

25000

H2 production (kg /h)

Base case Alternative 1: higher amount of ash Alternative 2: more C, H; less O, S; same N, ash

25000

Phas es leaving the s eparator (kg /h)

30000

20000 15000

10000

Liquid

Gas

Char

5000

6000

14000

Alternative 1: higher amount of ash

12000

Electrical power (kWe)


4000

Products ( kg/h)

800 600 400

3000 2000

0

bio-oil

Base case

5000


200

Aqueous fraction

Ash

Alternative 1: higher amount of ash

1000

0

0

Base case

1000

H2 from SCWR

H2 from WGS

H2 from SR

H2 total

Base case Alternative 1: higher amount of ash Alternative 2: more C, H; less O, S; same N, ash

10000 8000 6000 4000 2000

0 Biofuel

LPG

Gasolina

0

Diesel

Turbine - SCWR Steam turbine

PEMFC

Consumption

Net production

Fig. 8. Effect of the composition of MSW feed on the fast pyrolysis (FP) production, phase separation, H2 production, biofuel production and electrical power under the same operating conditions.

aqueous fraction phases decrease; similarly, the H2 production is reduced to 923 kg/h. In addition, the different alternatives considered for the process in the development of the system are summed up in the Appendix B.

operating pressure, dimensions, and so on) and the manufacturing material, as mentioned in the previous section. By considering the contingency costs as well as the fees and costs of the auxiliary facilities, a total installation cost of 137.07 MM€ was obtained, by using the method by Turton et al. [29] and the method by Peters and Timmerhaus [30], which result in very similar estimate of the total cost of the installation. Production or operating costs are divided into fixed costs and variable costs. Table 11 includes the capital costs and the operating costs. Cooling water is the only utility required, as steam and electricity are produced in situ. Normally, the cooling water is heated from approximately 25 °C to 40–45 °C, and then cooled by a cooling tower to be reused. In this process, cooling water also produces steam, i.e., it has a process use, so typical cost of cooling water should not be applied. Thus, a higher cost of 0.81 €/t was assumed [29], assuming deionized water that may be converted into steam. Likewise, treatment of wastewater cost is 48 € per 1000 m3 [29]. The catalyst for the SCWR reactor could be Ni/Al2O3-SiO2, as previously verified [38]. In the WGS at high temperature usually is an iron oxide-chromium oxide-based catalyst, operating at temperatures from 310 °C to 450 °C. The WGS at low temperature use a copper-based catalyst. Typical compositions include Cu, Zn, Cr and Al oxides, and recent catalysts can be operated at medium temperatures up to 300 °C [39]. For the HDO reactor, a conventional sulfided NiMo catalyst could be used although sulfided Ru/C catalyst might be used in a first stage of the hydrotreater to achieve stabilization [40]. All the catalysts used in the different reactors were considered [39] and compared [41], and the cost of most expensive one (Ru/C) was considered as catalyst cost, to be more restrictive or conservative in the estimates, in the sense they will be likely higher than the real values. In addition, a low average space velocity of 20 kgfeed kgcatalyst-1 h−1 was used [42]. Similarly, the cost of the PSA sorbent was estimated from [43] assuming six beds, a mass feed-to-sorbent load ratio of 0.3 and a cycle time of 5 min. The resulting annual production cost is 11.85 M€. The revenues come from the sale of electricity and biofuels, as well as from the gate fee. This latter revenue is got from the authorities for taking charge of the waste (MSW reject fraction) and its treatment. Instead of obtaining the break-even points of LPG, gasoline and diesel, the gate fee was calculated and the prices of fossil fuels without taxes were adopted for biofuels: 0.765 €/kg LPG, 0.706 €/kg gasoline and 0.679 €/kg diesel [44]. The criterion is quite strict as the taxes for biofuels are lower than those for fossil fuels and they might be even

4.3. Design and cost of units Purchase equipment costs along with the references used to estimate them and the main characteristics of each designed unit are shown in Table 8. Among all the units, those whose cost was estimated by CAPCOST are briefly described. Thus, multiple pipe (multitubular) heat exchangers were selected for those units with an exchange area greater than 10 m2 operating at high pressure due to their good efficiency. Shell-and-tube floating head heat exchangers were selected when the pressure is lower (below 150 bar) and the area is high (up to 1000 m2). Double-pipe exchangers were used for units with an exchange area lower than 10 m2. The material is carbon steel, stainless steel or nickel alloy depending on the temperature and pressure of the streams. Vertical flash drums and horizontal decanters were designed to provide a residence time of 10 min for the liquid phase and the diameters were calculated based on the vapor flow-rate and the settling velocity of the liquid droplets. Similarly, a typical value for the heightto-diameter ratio of 3.0 was set for the decanters. Their costs were also estimated by CAPCOST. A reciprocating pump with a dampener, made of AISI 316, was chosen to pump bio-oil aqueous phase up to 240 bar (P1) and 100 bar (P2). The other pumps for cooling water are centrifugal. Regarding the compressors, both centrifugal and rotary types, depending on the gas flow-rate to be compressed. Likewise, two units are necessary for the two turbines because the total electrical power is higher than the maximum value accepted by CAPCOST. The costs of these units were obtained by CAPCOST, as well as for the distillation columns designed as mentioned in Section 3.1.3. The cost estimation of the other units was estimated from open access reports (Section 4.4). 4.4. Economic analysis Tables 9 and 10 show the cost of different units for a given capacity, once corrected the values by applying the CEPCI index, from scale up or down (Eq. (2)) and estimated using the CAPCOST tool, where the inputs are the required characteristics of each equipment (heat transfer area, 1178



null. In the same way, the electricity price for non-householder sector in the European Union was used (0.12 €/kW h in 2017 [45]). The biofuels have similar technical specifications to those equivalent fuels reported in norms such as UNE-EN-ISO 12185 and ASTM D 4052. By carrying out a cash flow analysis, the minimum gate fee to make the system profitable was determined by setting a net present value (NPV) equal to zero at the end of the lifetime plant. The gate fee is 16.7 €/t, and the assumptions were: 15% rate of return (with 100% equity financing, i.e., no money is borrowed), plant life time of 20 years, operation of 8000 h per year, a tax rate of 30% on profit, and

an amortization (applicable to the cost of the base module) performed linearly during the first 10 years. Additionally, other assumptions were: a working capital of 5% of the total investment cost, and the payment of 50% of the total capital at the beginning and the rest at the end of the first year, just when the plant is started-up. The variation of the cumulative non-discounted and discounted (present cash flow) over the lifetime of the plant is shown in Fig. 9. At the beginning (year zero in which the plant is built), the cash flow is negative because of 50% of TIC is paid. At the end of the first year, the process is ready to start, but the rest of capital costs (50% TIC) and

Table 8 Specifications and main design characteristics of the individual process units. Code

Equipment and Specifications

Design characteristics(1)

P1 P2 P3-P8 COMP1 (PSA 1)

Pump - Efficiency: 0.8; Outlet pressure: 240 bar Pump - Efficiency: 0.8; Outlet pressure: 100 bar Pumps of water - Efficiency: 0.75; Outlet pressure: 1.1–10.0 bar Train of compression; Type: Isentropic; Isentropic efficiency: 0.76; Mechanical efficiency: 0.98; two intermediate coolers and one final to 35 °C; Outlet pressure: 25 bar (compression ratio per stage of 2.92) Trains of compression; Type: Isentropic; Isentropic efficiency: 0.76; Mechanical efficiency: 0.98; two intermediate coolers and one final to 35 °C; Outlet pressure: 25 bar (compression ratio per stage of 2.92) Trains of compression; Type: Isentropic; Isentropic efficiency: 0.76; Mechanical efficiency: 0.98; two intermediate coolers and one final to 35 °C; Outlet pressure: 25 bar (compression ratio per stage of 2.92) Trains of compression; Type: Isentropic; Isentropic efficiency: 0.76; Mechanical efficiency: 0.98; two intermediate coolers and one final to 35 °C; Outlet pressure: 15 bar (compression ratio per stage of 2.92) Trains of compression; Type: Isentropic; Isentropic efficiency: 0.76; Mechanical efficiency: 0.98; one intermediate cooler and one final to 35 °C; Outlet pressure: 5 bar (compression ratio per stage of 2.24) Compressor for H2 - Type: Isentropic; Isentropic Efficiency: 0.76; Mechanical efficiency: 0.98; one intermediate cooler and one final to 35 °C; Inlet/Outlet pressure: 25/100 bar (compression ratio per stage of 2.00) Compressor for LPG (C-4) - Type: Isentropic; Isentropic Efficiency: 0.75; Mechanical efficiency: 0.98; one intermediate cooler and one final to 35 °C; Inlet/Outlet pressure: 5/25 bar (compression ratio per stage of 2.24) Fan - Type: Isentropic; Isentropic Efficiency: 0.76; Mechanical efficiency: 0.98; Outlet pressure: 1.1 bar (char combustion) Fan - Type: Isentropic; Isentropic Efficiency: 0.76; Mechanical efficiency: 0.98; Outlet pressure: 1.1 bar (pyrolysis gas combustion) Fan - Type: Isentropic; Isentropic Efficiency: 0.76; Mechanical efficiency: 0.98; Outlet pressure: 1.1 bar (Air for PEMFC) Turbine - Type: Isentropic; Efficiency isentropic: 0.85; Inlet/Outlet pressure: 240/40 bar (downstream from the SCWR reactor) Turbine - Type: Isentropic; Efficiency isentropic: 0.85; Inlet/Outlet pressure: 10/1 bar (steam from cooling water) Heat exchanger - Hot stream outlet temperature: 50 °C Heat exchanger - Cold stream outlet temperature: 300 °C Heat exchanger - Cold stream outlet temperature: 500 °C Heat exchanger - Cold stream outlet temperature: 240 °C Heat exchanger - Hot stream outlet temperature: 250 °C Heat exchanger - Hot stream outlet temperature: 400 °C Heat exchanger - Cold stream outlet temperature: 383 °C Heat exchanger - Cold stream outlet temperature: 385 °C (supercritical) Heat exchanger - Hot stream outlet temperature: 190 °C Heat exchanger - Cold stream outlet temperature: 35 °C Heat exchanger – Cold stream vapor fraction up to 1.00 Heat exchanger: Tin)hot fluid – Tout)cold fluid = 10 °C Heat exchanger - Hot stream outlet temperature: 190 °C Heat exchanger - Hot stream outlet temperature: 35 °C Heat exchanger: Tin)hot fluid – Tout)cold fluid = 10 °C Heat exchanger - Cold stream outlet temperature: 300 °C Heat exchanger - Hot stream outlet temperature: 35 °C Heat exchanger - Hot stream outlet temperature: 130 °C Heat exchanger – 1st increase in the vapor fraction as much as possible (0.30) Heat exchanger – 2nd increase in the vapor fraction as much as possible (0.74) Heat exchanger - Hot stream outlet temperature: 300 °C Heat exchanger - Cold stream outlet temperature: 80 °C Heat exchanger - 1st condensation stage of the steam from turbine (0.30) Heat exchanger - Cold stream outlet temperature: 160 °C Heat exchanger – Hot stream vapor fraction: 1.00 (2nd condensation stage) Heat exchanger - Hot stream outlet temperature: 35 °C Heat exchanger - Hot stream outlet temperature: 35 °C Gas-liquid separators (Flash); Cylindrical, vertical; Temperature: 35 °C; SEP 1 for pyrolysis gas and vapor (1 bar, 50 °C); SEP2 for SCWR-WGS product (40 bar); SEP3 for HDO product (100 bar); SEP4 for SR product (25 bar, variable); SEP5 for LPG product (25 bar)

SS; reciprocating SS; reciprocating CS; centrifugal CS; centrifugal

348.0 kWe 57.4 kWe 35.7 kWe 1,928.7 kWe

CS; centrifugal

1,318.5 kWe

COMP2 (PSA 2) COMP3 (PSA 3) COMP4 (PSA 4) COMP5 (PSA 6) COMP6 COMP7 FAN1 FAN2 FAN3 TURB TURB2 HE01 HE02 HE03 HE04 HE05 HE06 HE07 HE08 HE09 HE10 HE11 HE12 HE13 HE14 HE15 HE16 HE17 HE18 HE19 HE20 HE21 HE22 HE23 HE24 HE25 HE-26 HE-27 FLASH-1 FLASH-2 FLASH-3 FLASH-4 FLASH-5

Capacity

CS; rotary

287.7 kWe

CS; centrifugal

816.7 kWe

CS; rotary

50.1 kWe

CS; centrifugal

2,278.0 kWe

CS; rotary

52.5 kWe

CS; centrifugal radial

36.1 m3/s


6.7 m3/s


6.8 m3/s

Nickel alloy; axial expander

5,517.9 kWe(2)

SS; axial steam turbine

7,851.6 kWe(2)

SS-CS; floating head SS-CS; floating head SS-SS; double pipe SS-SS; multiple pipe SS-SS; multiple pipe SS-SS; multiple pipe SS-SS; multiple pipe Ni-Ni; double pipe SS-CS; floating head SS-CS; floating head SS-CS; floating head SS-CS; floating head SS-CS; floating head SS-CS; floating head SS-CS; multiple pipe SS-SS; multiple pipe SS-CS; floating head SS-SS; double pipe SS-CS; floating head SS-CS; floating head CS-CS; multiple pipe SS-SS; double pipe CS-CS; floating head CS-CS; double pipe CS-CS; floating head SS-CS; floating head SS-CS; floating head 2.02 m ID; 5.61 m height; SS 1.23 m ID; 5.96 m height; SS 1.45 m ID; 4.71 m height; SS 1.41 m ID; 4.36 m height; SS 0.35 m ID; 2.34 m height; SS

2,161.0 m2(3) 12.0 m2 2.1 m2 223.0 m2(3) 98.4 m2 22.1 m2 19.2 m2 2.2 m2 216.7 m2 277.2 m2 12.0 m2 291.4 m2 21.4 m2 587.2 m2 542.0 m2 0.5 m2 157.2 m2 5.9 m2 29.0 m2 74.6 m2 48.3 m2 0.4 m2 742.0 m2 5.0 m2 1,147.0 m2(2) 17.6 m2 18.5 m2 17.9 m3 7.1 m3 7.8 m3 6.8 m3 0.2 m3

(continued on next page) 1179



Table 8 (continued) Code

Equipment and Specifications

Design characteristics(1)

DECANT1

Liquid-liquid separator for pyrolysis product (Decanter); Temperature: 39.2 °C.; Pressure drop: 0.1 bar; Cylindrical, horizontal (1 bar) Liquid-liquid separator for HDO product (Decanter); Temperature: 35 °C.; Pressure drop: 0.1 bar; Cylindrical, horizontal (25 bar)

2.15 m ID; 6.46 m height; SS

23.47 m3

1.38 m ID; 4.13 m height; SS

6.16 m3

Fast pyrolysis reactor (RYield, Operating temperature and pressure: 500 °C, 1 bar); CFB reactor, using inert solid as a medium of heat transfer (fluidified by 64,784 kg/h fluidization gas) RYield + RGibbs; Combustion of everything able to be oxidized; TRYield: 500 °C @ 1 bar; TRGibbs: energy balance (1150 °C); Surplus heat flow to FP reactor and for heating the fluidization gas in the FP-CFB reactor; flue gas is used in the SR reactor Supercritical Water Reforming (RGibbs, Operating temperature: 800 °C) + Furnace-combustor (RStoic, combustion of everything able to be oxidized; 1000 °C @ 1 bar, surplus heat flow to SCWR) High-temperature Water Gas-Shift reactor; REquil; Operating temperature: 400 °C; CO + H2O ↔ CO2 + H2 (adiabatic); Outlet temperature: 425 °C Low-temperature Water Gas-Shift reactor; REquil; Operating temperature: 250 °C; CO + H2O ↔ CO2 + H2 (adiabatic); Outlet temperature: 252 °C Steam reforming reactor; RGibbs; Isothermal reactor: 800 °C; Operating pressure: 25 bar (changed in the sensitivity analysis); CH4 + H2O ↔ 3H2 + CO (& WGS at equilibrium: CO + H2O ↔ CO2 + H2) Hydrodeoxygenation reactor (RYield); Operating temperature: 300 °C; Pressure: 100 bar; Products and enthalpy obtained by calculator; 57.3 wt% H2O; 9.2 wt% H2; 4.4 wt% CO2; 7.6 wt% C1-C5; 11.6 wt% C6-C10; 9.2 wt% C12-C18 Stoichiometric reactor; Heat + Work to SPL-PEM (efficiency: 45%); H2 + 0.5 O2 → H2O; Isothermal Reactor: 80 °C First PSA system: It removes most the H2 (95%) from the other gases. Outlet pressure: 25 bar (pure H2 stream), 1.01 bar (the rest of gases) Second PSA system: CO + N2-rich stream (98% total CO and 100% total N2 at the top): 14.0 wt% feed; 99% total CO2 and 99% total CH4 at the bottom: 86.0 wt% feed. Outlet pressure: 25.0/1.01 bar (top/bottom) Third PSA system: CO-rich stream (93.7% CO, 5.5% CO2 and 0.6% CH4) at the bottom: 91.6 wt% feed; 17.6% N2 and 82.3% H2 at the top: 8.4 wt% feed. Outlet pressure: 25.0/1.01 bar (top/bottom) Fourth PSA system: It removes most the CH4 (85%) from the other gases, and 10% of CO2 entering the first PSA system; pure CO2 stream (bottom): 81% feed (89.1% total CO2, 8.9% total CH4); CH4 (top): 19% feed. Outlet pressure: 25.0/1.01 bar (top/bottom) Fifth PSA system: It removes all the N2 (85%) from the H2. There is no need of further compression like in the other PSA systems. Sixth PSA system: It removes the CO2 from the other gases (C2-C6). The gas used for regeneration is CH4 instead of CO2 to get a better performance of the distillation system. Thus, 17% of CH4 removed in the fourth PSA system is derived to the sixth PSA system. Outlet pressure: 15.0/1.01 bar (top/bottom) Distillation column of 60 sieve plates each one in series (30 ideal stages; stage efficiency: 50%); To remove gases and low boilers (C2-C5); 0.56 m ID; 40 m height; 0.6 spacing; down-comer area: 10% column area; carbon steel. Temperature at the top: 40.7 °C, Temperature at the bottom: 285.4 °C; 25 bar Distillation column of 40 sieve plates each one in series (20 ideal stages; stage efficiency: 50%); To separate low boilers (C2-C6) with 16 wt% C7+; 0.28 m ID; 28 m height; 0.6 spacing; down-comer area: 10% column area; carbon steel. Temperature at the top: 128.9 °C, Temperature at the bottom: 180.8 °C; 5 bar Distillation tower of 100 sieve plates each one in series (50 ideal stages; stage efficiency: 50%); Gasoline at the top (C6-C10) with 0.2 wt% C4; 10.1 wt% C5; 1.8 wt% C12+; Diesel at the bottom (C12-C18) with 10.5 wt% C10; 0.55 m ID; 64 m height; 0.6 spacing; down-comer area: 10% column area; carbon steel. Temperature at the top: 129.7 °C, Temperature at the bottom: 306.3 °C; 3 bar Distillation tower of 22 sieve plates each one in series (11 ideal stages; stage efficiency: 50%); LPG at the top (C2-C5) with traces C4+; 0.55 m ID; C5+ at the bottom with 1.9 wt% C4; 17 m height; 0.6 spacing; down-comer area: 10% column area; carbon steel. Temperature at the top: 129.7 °C, Temperature at the bottom: 306.3 °C; 5 bar

44,011 kg/h

DECANT2

FP reactor Char boiler SCWR + FURNACE HWGS LWGS SR reactor HDO reactor PEM fuel cell PSA1 PSA2 PSA3 PSA4 PSA5 PSA6 C-1 C-2 C-3 C-4

Capacity

34.2 MWth 31,982 kg/h 8.91 kg/s 8.91 kg/s 4.48 kg/s 6.34 kg/s 5,131.1 kW 49,074 Nm3/h 11,794 Nm3/h 3,745 Nm3/h 7,832 Nm3/h 1,961 Nm3/h 6,131 Nm3/h 9.85 m3 1.72 m3 15.20 m3 4.04 m3

(1) Materials: CS: carbon steel; SS: Stainless steel (AISI 316); Ni: nickel alloy (such as Inconel 600 or 625). (2) Two units. (3) Three units are necessary to achieve this value.

Table 9 Specifications and main design characteristics of the individual process units. Unit Pyrolysis reactor Char boiler SCWR reactor + furnace WGS (two stages) HDO (estimated as an HDS) PSA1 system PSA2 system PSA3 system PSA4 system PSA5 system PSA6 system SR reactor H2S removal system Compressor for H2 Steam turbine PEM fuel cell Refrigeration system

Base cost (CB)

Reference currency and year

Base capacity (AB)

5.98 1.97 2.54 2.78 13.19 0.84

M$2014 M$2014 M€2013 M$2014 k$2002 M€2015

14.3 20 3774.9 150 18.9 944

71.87 0.36 0.49 0.34 53.0 1.06

k$2002 M€2016 M$2014 M$2014 $2015 M$2014

79.06 1000 413 10.5 1 500

Scale factor (n)

t/h inlet MW kg/h kg/s kg/h feed Nm3/h feed

0.7 0.83 0.60 0.67 0.70 0.55

kg/h feed Nm3/h kW MW kW kW

0.67 0.60 0.68 0.44 1.00 0.67

1180

Capacity (AA) 44.01 34.21 32,066 8.88 22,807 49,074 11,974 3,745 7,832 1,961 6,131 16,122 49,074 1,502 7.85 5,131 87.2

t/h entrada MW kg/h kg/s kg/h feed Nm3/h feed Nm3/h feed Nm3/h feed Nm3/h feed Nm3/h feed Nm3/h feed kg/h feed Nm3/h kW MW kW kW

Cost (CA) (M€2017)

Ref.

13.32 3.12 9.42 0.81 2.29 6.02 2.75 1.46 2.19 1.02 1.92 3.50 3.28 1.14 0.26 0.23 0.31

[31] [32] [32] [32] [33] [34]

[33] [35] [31] [31] [36] [31]



parameters in a project, since the method of cost estimation followed has an expected accuracy of roughly ± 30%. Thus, a sensitivity analysis was carried out to assess the effect of uncertainty in TIC ( ± 30%) from the base scenario, using steps of 10% around base scenario (0). Fig. 10 shows how the gate fee changes with the TIC, so if this latter decreased 19% the gate fee would be null; but if the TIC increased to 178 MM€ the gate fee would be still lower (43.3 €/t) than normal values, as above mentioned. Other variations such as higher biofuels or electricity prices will lead to a better scenario, i.e., to a negative gate fees, which means a decrease in the selling price of biofuels if no gate fee was considered and, hence, a more profitable plant.

Table 10 Purchase cost of main units using CAPCOST. Unit

Cost (M€2017)

Heat exchangers Distillation columns Distillation condensers Distillation reboilers Compressors (PSA/LPG) Fans

7.71 3.63 0.19 0.40 5.55 0.06

Unit

Cost (M€2017)

Pump P1 Pump P2 Pumps P3-P8 Flash drums Decanters Cyclones*

0.16 0.10 0.07 0.77 0.26 0.25

* Estimated by the website of Peters and Timmerhaus [37].

working capital (5% TIC) are paid, so this is the point with the most unfavorable cash flow. From the second year onwards, money from sales (biofuels and electricity) is received because the production begins, as well as the gate fee, and the investment starts to be recovered with an internal rate return of 15%. Regarding the number of years to recover the initial investment of the project (discounting the land and working capital), the simple (non-discounted) and discounted payback period are 6.5 and 15.5 years, respectively, from the beginning. As the payback period is lower than lifetime of the project, this may be profitable. As abovementioned, the gate fee is 16.7 €/t, which is lower than those found in the literature [46], usually between 30 and 80 €/t, depending on the type of location and the type of installation (landfill, incineration or composting). Other values found [47] place the average gate fee for landfills in Europe at 80 €/t. In a recent publication [48], the gate fee necessary to obtain a rate of return of 15% was calculated for processes of waste gasification and incineration. In the case of incineration, the only one among those with a plant capacity similar to that used in this work, resulted in a gate fee of about 80 €/t. In the case of gasification, the gate fee values range from 50 to 130 €/t, decreasing as the capacity of the system increases. The total investment cost (TIC) is one of the most influential

5. Conclusions In this study, a conceptual design of a process to valorize the MSW reject fraction was developed. Thus, by fast pyrolysis a bio-oil is obtained, which is separated into an aqueous fraction phase that is reformed under supercritical conditions to produce H2 to be used in an HDO unit to upgrade the oil fraction phase. A high fraction of the carbon fed to the process may be converted into biofuels and net electrical power is obtained. This way, for a feeding of 50 t/h of MSW, a net electric power of 10.65 MWe is generated and 5.2 t/h biofuels are produced. A complete economic assessment resulted in a very low gate fee of 16.7 €/t, using the same industrial selling prices of fossil fuels and electricity in a full plant, assuming 15% rate of return (with 100% equity financing) and plant life time of 20 years, so the proposed process may be feasible. In fact, the concept seems to be economic because the process is fully integrated and energy self-sufficient, and the design has been performed to get maximum production of biofuels and electricity trying to take advantage of all the streams and flows.

Table 11 Estimated investment costs and operating costs assuming 8000 h per year. Capital cost (k€2017) -method 1-1 CBM (bare module cost) Contingency and fees Auxiliary facilities Grassroots costs

81,441.27 14,659.43 40,720.64 136,821.33

Fixed operating costs

Purchase equipment and installation (PECI) Instrumentation and control Piping Electrical systems Services (facilities) Land Building (including services) Yard improvements Total Direct Costs (TDC) INDIRECT COST Legal expenses Contractoŕs fee Startup assistance EPC constructor contingency Total indirect costs Total investment cost (TIC)

%PECI 100 8 15 5 5 2 8 10 153 %TDC 2 3 1 4 10

(k€2017) -method 2-2 81,441.27 6,515.30 12,216.19 4,072.06 4,072.06 1,628.83 6,515.30 8,144.13 124,605.14 2,492.10 3,738.15 1,246.05 4,984.21 12,460.51 137,065.66

Variable operating costs

Concept

%TDC

€2017/year

Labor Maintenance General expenses Management and operation services Cost of marketing, logistic and others Insurance Total

1.55 1.50 3.07 0.44 1.32 0.50 8.38

1,931,379.72 1,869,077.15 3,825,377.91 548,262.63 1,644,787.89 623,025.72 10,441,911.03

Total (€2017/year)

Direct cost

Concept Catalyst HDO (€/kg) Catalyst SR (€/kg) Catalyst 2-WGS (€/kg) Catalyst SCWR (€/kg) Sorbent 6-PSA (kg/h) MSW (€/kg) Process/cooling water (€/t) Waste water treatment (€/m3) Total

(1) The method by Turton et al. [29] and (2) The method by Peters and Timmerhaus [30]. 1181

Unit cost

€2017/year

106.92

122,958.00 144,342.00 342,144.00 122,958.00 387,090.00 0.00 3,929,310.00 232,848.72 5,281,650.00 15,723,561.03

60.59 0 0.81 0.0048



360 320 280 240 200

M€

160 120 80 40 0 -40

Non-discounted cash flow Discounted cash flow

-80

PB for investment recovering

-120 -160

0

1

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

Lifetime (year) Fig. 9. Cumulative non-discounted and discounted cash flow diagrams over the lifetime, by considering 100% equity financing.

and no cold stream comes into it. As a result, the design rules for networks of heat exchangers from the pinch cannot be used, and the maximum energy recovery between the streams must be analysed in other way. The strategy followed is that streams close to each other at high thermal levels are the first to contact with each other (the high-temperature hot streams warm up the high-temperature cold streams), and streams at low thermal levels are the last to contact with each other (the low-temperature hot streams heat the low-temperature cold streams). This way, the feedings to the HDO and SR units are first preheated with the products leaving those reactors. For the SCWR reactor, the feeding is first preheated with the output of the second WGS stage, then with the product of the first WGS stage, and afterward with the output stream of SCWR, once expanded to 40 bar. The energy available in the reaction products is not enough to heat the reactants to the required temperature in HDO, SR and SCWR reactors. For this reason, flue gases generated in the two furnaces of the overall process are used to complete the heating of the reactants. Regarding the heating, the flue gas generated at 1000 °C in the furnace associated with the SCWR reactor, where the pyrolysis gas stream is burnt, is firstly used to preheat the pyrolysis gases, thus increasing the fraction of the combustion heat added to the SCWR reactor. Then, after preheating of the feeding to SCWR reactor until reaching the supercritical state, it preheats successively the air used in the furnace, the hydrogen entering the fuel cell and the cooling water used for district heating. Likewise, the flue gas generated in the second furnace where the pyrolysis char is burnt, firstly, provides the heat needed to reach and maintain the reaction temperature (800 °C) in the SR reactor. Then, it is used in the reboilers of the four distillation columns; after, this stream heats up the HDO feeding and preheats the feeding to the SR reactor. Afterward, the flue gas is used to increase the steam fraction of the cooling water and to superheat the generated steam, thus increasing the electric power generated in the steam turbine. Finally, it is used to dry the MSW fed to the system before being emitted to atmosphere. With respect to the cooling, as a temperature of 35 °C must be reached at different points for condensation and water separation,

40 35 30

Gate fee (€/t)

25 20 15 10 5 0 -5 -10 -15 -20 -30%

-20%

-10%

0%

10%

20%

30%

Uncertainty in investement costs Fig. 10. Sensitivity analysis of the gate fee versus the uncertainty of the total investment costs.

Appendix A. Heat integration An energy and material integration of the process as complete as possible must be carried out in order to achieve a very high process efficiency. In addition, as an imposed constraint, the process must be energy self-sufficient, without using external fuels for heating or electricity from electrical grids at steady state conditions. Likewise, the process is aimed at reducing as much as possible the consumption of utilities, such as cooling water. To carry out the energy integration, the design of the heat exchangers network is usually carried out on a pinch analysis. The pinch is the point where the temperature difference between the hot composite curve and the cold composite curve has the minimum allowable value. In this process, the energy difference between the streams to be cooled and heated causes the pinch to be located at the highest temperature point. Moreover, only one hot stream comes out of the pinch 1182



different cooling water streams at 25 °C must be used. If the water pressure is 10 bar, medium pressure steam can be generated. One of those streams allows the cooling of the stream leaving the second stage of WGS. Another stream cools the stream leaving the SR reactor in a similar way. The third cooling water stream is used for cooling the stream leaving the HDO reactor and the HDO reactor itself. These three streams join in one at 180 °C where the steam fraction is 0.3, which is increased to 0.74 by using the residual thermal energy of the flue gas from the char combustion, as above mentioned. The saturated steam is separated by a flash unit and further superheated to 300 °C by exchanging with the former flue gas stream before entering the steam turbine, where is expanded to generate 7.85 MWe. In addition to the three previous cooling water streams, three water streams at atmospheric pressure are used to cool the condensed pyrolysis liquid, once the quench has been produced, and in the two condensers after the steam turbine. The other streams are used in the distillation condensers (four), as well as for conditioning the gasoline and diesel. As the other streams, these two new water streams are also used as heating water in the installation. Finally, a refrigerated cooling water at 5 °C is needed to obtain the LPG at liquid state in the condenser of the distillation column C-4. The water is after used in the second stage of steam condensation (HE-25).

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Appendix B. Process alternatives During the development of the system, different alternatives of the process were considered, but finally disregarded, as explained below. First, the compression of the pyrolysis gas just at the outlet of the fast pyrolysis to treat them in the PSA system would allow the separation of the H2 present. This way, the hydrocarbon stream leaving the PSA fed by the pyrolysis gas stream, unreacted gas from the SCWR and gas formed in the HDO, would be burnt in the furnace to provide the heat needed for the SCWR. However, the compression cost of the pyrolysis gas up to the pressure required in the PSA is much higher than the extra electric generation obtained by the separated H2. In addition, to get the heat required in the SCWR, the flow-rate processed in the SR would have to be reduced, thus decreasing H2 production at this point. Second, the gas leaving the HDO process at 100 bar could be reduced to 25 bar, before entering the PSA system, using a turbine to generate extra electric power. However, the power generated would be small, due to the relatively low flow-rate of this stream. As a result, the extra complexity associated with an additional turbine would not be compensated. Third, if the pressure at the outlet of the turbine located downstream from the SCWR reactor were 100 bar, instead of 40 bar, the H2 compression would be avoided before entering the HDO process. However, the lower power produced in the expander is higher than the power consumed in the H2 compression. In addition, when increasing the pressure, the H2 yield in the SR reactor is lowered, so a greater recirculation of unconverted reactants would be necessary. References [1] “Database - Eurostat.” Available: http://ec.europa.eu/eurostat/data/database [accessed 20.07.18]. [2] Abeliotis K. Life cycle assessment in municipal solid waste management. Kumar Sunil, editor. Integrated waste management, vol. 1. InTech; 2011. [3] Montejo C, Costa C, Ramos P, Márquez MC. Analysis and comparison of municipal solid waste and reject fraction as fuels for incineration plants. Appl Therm Eng 2011;31:2135–40. [4] Brás I, Silva ME, Lobo G, Cordeiro A, Faria M, Teixeira L, et al. Refuse derived fuel from municipal solid waste rejected fractions-a case study. Energy Procedia 2017;120:349–56. [5] Chen D, Yin L, Wang H, He P. Pyrolysis technologies for municipal solid waste: a review. Waste Manage 2014;34:2466–86. [6] Czajczynska D, Anguilano L, Ghazal H, Krzyzynska R, Reynolds AJ, Spencer N, et al. Potential of pyrolysis processes in the waste management sector. Therm Sci Eng Progr 2017;3:171–97. [7] Bridgewater A. Biomass fast pyrolysis. Therm Sci 2004;8:21–50.

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