Techno-economic analysis of a co-electrolysis-based synthesis process for the production of hydrocarbons

Applied Energy 215 (2018) 309–320

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Applied Energy journal homepage: www.elsevier.com/locate/apenergy

Techno-economic analysis of a co-electrolysis-based synthesis process for the production of hydrocarbons

T

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Gregor Herz , Erik Reichelt, Matthias Jahn Fraunhofer IKTS, Fraunhofer Institute for Ceramic Technologies and Systems, Winterbergstraße 28, 01277 Dresden, Germany

H I G H L I G H T S coupled process of co-electrolysis and Fischer-Tropsch synthesis was developed. • AHeat integration allows for an energetic efficiency > 60%. • Economic can be improved when valuable products are focused. • Availabilityfeasibility and price of renewable energy are crucial for industrial application. •

A R T I C L E I N F O

A B S T R A C T

Keywords: Co-electrolysis Fischer-Tropsch synthesis Process modeling Techno-economic analysis Power-to-X

The paper is focused on the development and techno-economic assessment of a sustainable process for the production of valuable hydrocarbons from CO2 and H2O. With help of process modeling tools an advantageous and highly integrated process design is identified. The application of co-electrolysis for direct syngas production as well as the implementation of advanced heat integration concepts allow for an energetic efficiency of ηen > 0.6. Additionally, the amount of reactors and heat exchangers in the proposed process is reduced in comparison to other Power-to-X concepts. The high efficiency and the focus on valuable products like waxes is also beneficial for the economic feasibility of the process. The implications of product value as well as availability and costs of renewable electricity are discussed in the context of a potential market entry of Power-to-X technologies.

1. Introduction The utilization of renewable energy for the sustainable production of chemical products from CO2 and H2O is of increasing interest worldwide. Different process concepts based on the coupling of electrolysis and a chemical synthesis step are summarized under terms like Power-to-Gas, Power-to-Fuel or Power-to-Chemicals. The variety of electrolysis technologies (alkaline, polymer electrolyte membrane and solid oxide electrolysis) and potential synthesis steps (e.g. methanation, methanol synthesis, Fischer-Tropsch synthesis) allow for several different process concepts. However, so far the technical demonstration of such Power-to-X concepts is particularly focused on the production of methane. Here several demonstration plants exist [1]. Plants for the production of liquid fuels are only operated by Carbon Recycling International [2] and Sunfire [3]. Especially the production of liquid fuels is believed to play a crucial role in a future energy and transport system [4]. However, for the mass market of transportation, production costs are of high importance. As

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recent studies show, the production costs for so-called electrofuels are considerably higher than the production costs for fossil fuels [5]. For the necessary reduction of production costs, an increase in efficiency of the Power-to-X technologies is of vital importance. As the market entry of Power-to-X technologies on a mass market like transportation is expected to be more difficult, first applications should be focused on niche markets, where advantageous products are able to achieve high prices. Potentially interesting products are waxes produced via Fischer-Tropsch synthesis. Due to the absence of aromatic and polycyclic aromatic compounds as well as sulfurous substances, Fischer-Tropsch waxes are especially feasible for cosmetic applications [6]. The increasing demand for waxes [7] and the sustainable production makes this synthesis route an attractive option for a first market entry of Power-to-X technologies. Established electrolysis technologies like alkaline and PEM (polymer electrolyte membrane) electrolysis are only capable of converting H2O to H2 [8,9]. The direct conversion of H2 and CO2 to hydrocarbons is limited to only few reactions (e.g. methanation, methanol

Corresponding author. E-mail address: [email protected] (G. Herz).

https://doi.org/10.1016/j.apenergy.2018.02.007 Received 31 August 2017; Received in revised form 1 February 2018; Accepted 2 February 2018 Available online 20 February 2018 0306-2619/ © 2018 Elsevier Ltd. All rights reserved.

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G. Herz et al.

for internal reforming in comparison to the case of reverse water-gas shift reaction for tail gas conversion. In both cases, a significant increase in efficiency compared to a once-through process was observed. Even though the process concept proposed in Ref. [19] allows for a high overall efficiency, it would have to be altered in order to be technically implemented. For example the necessity for a purge stream to avoid the accumulation of inert components was not considered. However, the proposed concept highlights the importance of by-product recirculation for process efficiency. The potential of heat integration via utilization of the purge gas was evaluated by Becker et al. [20]. A simplified flowsheet of the proposed process concept is depicted in Fig. 2b. In this case, by-products and unconverted syngas are combusted in order to supply thermal energy to the syngas production step. The purge gas is combusted with anode off-gas in order to reach high exhaust temperatures, which significantly improves heat integration. Furthermore, thermal losses caused by inert nitrogen can be minimized. The hot exhaust gas is first used to pre-heat the inlet streams of cathode and anode side of the SOEC. After exiting the preheating steps, the exhaust gas can be used to pre-heat carbon dioxide, evaporate feed water and supply heat to syngas pre-heating, product upgrading and an electric generator. In this process concept, the possibility of separating the hydrogen from the tail gas was also introduced (not included in Fig. 2b in order to aid comparability). A fraction of this hydrogen is fed to the SOEC to avoid degradation [21,22], while the rest is considered as a marketable product. This separation step significantly increases overall process complexity and is therefore not reasonable for a technical realization. A process concept proposed by König et al. [23] incorporates both mentioned possibilities of tail gas utilization but utilizes a water electrolysis and a reverse water-gas shift step. A part of the tail gas is recycled and again converted into syngas in the reverse water-gas shift step while the rest is combusted to supply heat to the rWGS reactor. The ratio between these two streams is adjusted, so that the heat demand of the syngas production step is met. While this approach allows for a very flexible process management, not all advantages of high temperature electrolysis were considered, such as co-electrolysis or internal reforming. Additionally, in this concept the heat of the exothermic Fischer-Tropsch synthesis is not utilized. The main characteristics of the aforementioned process concepts are listed in Table 1. This summary illustrates the variety of different implementation methods for coupled processes of high temperature electrolysis and Fischer-Tropsch synthesis. In the process concept depicted in Fig. 3, the aforementioned possibilities for tail gas utilization (short recycle (a), long recycle (b) and tail gas combustion (c)) are combined. Additionally, a concept for the utilization of the heat produced in the exothermic synthesis step introduced by Verdegaal et al. [3] was included. The heat is used to evaporate the feed water supplied to the process. This reduces the thermal energy demand for reactant conditioning significantly, as shown in Section 3.1.2. This form of heat integration has so far not been implemented into a model of a Power-toX process. The process scheme in Fig. 3 is the base for the process design study in this paper. With help of process modeling an advantageous process design for the highly efficient production of waxes from renewable energy is developed.

synthesis) [10]. For more complex and more valuable products, syngas is necessary as feed. In this case, an additional reverse water-gas shift reaction step is necessary for syngas production [3]. A highly efficient alternative is the application of solid oxide electrolysis cells (SOEC). This technology allows for the direct production of syngas from CO2 and H2O in a so-called co-electrolysis. The source of the CO2 used as feed is important for Power-to-X processes. Potential point sources of CO2 are industrial emitters (e.g. steel mills, cement plants) and power plants. The separation of CO2 from the diluted flue gas streams of these emitters is costly. Expected costs for the separation depend on the emitter and differ considerably in the range of approximately 25–300 €/tCO2 [5,11,12]. The utilization of CO2 from the mentioned emitters would allow for a reduction of crude oil usage and would therefore lead to an emission reduction. However, to achieve the ambitious CO2 reduction goals of the climate summit in Paris in 2015, the utilization of fossil fuels must be significantly reduced. The increasing share of renewables in electricity generation [13] and novel production processes, e.g. in the steel industry [14,15] will lead to a reduction in available CO2 point sources in the future. Therefore, CO2 capture from air might be an interesting alternative in the long-term. First industrial solutions are available [16]. However, the costs are so far rather high with reported values in the range of 100–1000 €/tCO2 [5,17,18]. The above considerations show that economic feasibility of Powerto-X plants is highly influenced by several factors. Some of these factors can be influenced by legislation, e.g. by remunerations for avoided CO2 emissions or reduced electricity costs for Power-to-X plants. However, in order to allow for a future application of such processes, concepts with optimized energy and cost efficiency have to be found. Therefore, in this paper, a novel highly integrated process for the co-electrolysisbased production of chemicals is proposed. The efficiency of the process is calculated with help of process modeling tools and compared to other concepts. Based on these results, the economic feasibility of the process is evaluated. The influence of availability, electricity price and product fraction is considered and discussed. The results not only allow for an evaluation of Power-to-X processes, but also give an indication which potential tools might support the market entry of Power-to-X technologies. 2. Process and model development 2.1. Process scheme Generally, a Power-to-X process can be divided into three main process steps: syngas production, synthesis step and product separation. A simplified process scheme is depicted in Fig. 1. There are various different means of implementing these process steps. They are discussed in Section 2.2. This section focuses on general concepts for recirculation and heat integration. Power-to-X processes have been studied in the literature by several authors and it was found that by-product utilization is crucial in order to reach an efficiency competitive with other sustainable hydrocarbon synthesis processes. In a process concept proposed by Cinti et al. [19], two different means of by-product utilization were implemented (Fig. 2a). In this process scheme all by-products (water, tail gas) are recycled via a long recycle into the syngas production step and reformed (as opposed to a short recycle into the synthesis step). The conversion of by-products via internal reforming as well as in a reverse water-gas shift step were considered. The simulation results showed a slightly higher efficiency

electricity reactants

syngas production

synthesis

2.2. Modeling The process was modeled using the flowsheet based simulation

product separation

310

Fig. 1. Simplified process scheme of a Power-to-X process.

products

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tail gas

FischerTropsch reactor

syngas production

H2O, CO2

water

a)

products

tail gas

exhaust

combustion air

FischerTropsch reactor

syngas production

H2O, CO2

products water

b) Fig. 2. Simplified flowsheet according to Cinti et al. [19] (a) and to Becker et al. [20] (b).

studied and compared. Since the process consists of several pressurized or high temperature steps, the property model PSRK was chosen for all blocks except condensation steps. The PSRK model yielded inaccurate results for liquid/ gaseous equilibria at low temperatures. For these process steps, the RKSOAVE model proved to be much more reliable. The component data is derived from the APV88 PURE32 database in both cases. While alkanes up to a chain length of 30 carbon atoms were available in the database, longer chained components were not. To include the product distribution that can be obtained in a cobalt-based Fischer-Tropsch synthesis, components up to a chain length of 70 carbon atoms were added using the ‘Molecular Structure’ input form. The main process steps syngas production and synthesis were embedded in hierarchy blocks for convenient replacement. The product separation was modeled with simple condensation blocks, because the product can be upgraded on the demand side of the market with stateof-the-art processes, as it is proposed for small-scale Fischer-Tropsch

Table 1 Summary of the main characteristics of the Power-to-X processes proposed in literature. Reference

Syngas production

Synthesis

Tail gas utilization

Becker et al. [20] Cinti et al. [19]

Co-electrolysis

Fischer-Tropsch (cobalt catalyst) Fischer-Tropsch (cobalt catalyst)

Combustion

König et al. [23]

Co-electrolysis rWGS Internal Reforming Water electrolysis and reverse water-gas shift

• •

Fischer-Tropsch (iron catalyst)

Long recycle

Short recycle Long recycle Combustion

package Aspen Plus. The flowsheet was structured in a modular fashion that allows for the convenient substitution of the process steps while ensuring constant boundary conditions. This enabled various possibilities for the integration of the SOEC in a Power-to-X process to be

tail gas

c)

exhaust

combustion air

a) b) CO 2

cond1 syngas production

dry syngas

FischerTropsch reactor

cond2

cond3

waxes

H2Og

water evaporation

liquid products

H2Ol Fig. 3. Simplified process scheme applied for the process design study in this paper. Depicted are the three possibilities for tail gas utilization: (a) short recycle, (b) long recycle and (c) tail gas combustion.

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air

enriched air

data (e.g. [20]). Due to these models being specific for a chosen stack design and performance, a general evaluation of several process routes is difficult. Additionally, there are no models available that include internal reforming. Due to this and the fact that the results presented here should not only be applicable to one certain cell design, a shortcut approach was used in this work. Since the cell voltage Ucell is a function of the current density i, it was assumed that by varying the cell area and thereby the current density, a predefined value for the cell voltage can be achieved. In this work thermoneutral conditions for the operating point of the SOEC were considered, which means that the electric energy demand as well as the thermal energy demand of the electrolysis reactions are covered by electricity:

anode cathode

feed syngas RGIBBS

RSTOIC

RGIBBS

Fig. 4. Aspen Plus flowsheet for the SOEC according to Cinti et al. [19].

processes [19,24,25].

Pel = ΔḢ .

2.2.1. Syngas production The syngas necessary for the Fischer-Tropsch synthesis is supplied by an SOEC. A significant advantage of using an SOEC compared to low temperature electrolyzers is the possibility of converting water and carbon dioxide simultaneously. The according reactions are

The thermal energy for the electrolysis reaction is provided by electricity being converted into heat due to losses in the cell, while the thermal energy for the reforming reactions is to be supplied externally. A comparable approach was applied by Cinti et al. [19]. It allows for the generation of reliable data for process evaluation. Based on the modeling results, it is possible to calculate the necessary cell area for a specific SOEC stack, if a suitable SOEC model or a polarization curve is available [19]. Under the assumption that the influence of the reforming reactions on the cell voltage is negligible, the cell voltage is a function of the energy demand of the electrolysis reaction and the molar flux of oxygen ions through the electrolyte:

H2 O⇄ H2 +

1 O2, Δh0 = +285.8 kJ/mol 2

(1)

CO2 ⇄ CO +

1 O2 . 2

(2)

Δh0 = +283.2 kJ/mol

In order to drive the reactions, an overall energy demand ΔRh has to be met. In case of a reversible process, this energy demand consists of a thermal energy demand TΔs and an electric energy demand ΔRg. While the overall energy demand ΔRh is not significantly influenced by temperature, the ratio between thermal and electric energy demand depends on the reaction. Generally, for both reactions the electric energy demand decreases with increasing temperature [8]. The mechanisms of the electrochemical conversion of water and carbon dioxide in an SOEC are still subject of intense research. While some models are based on the assumption, that the water-gas shift reaction

CO + H2 O⇄ CO2 + H2, Δh0 = −41.2 kJ/mol

Ucell =

Δh0 = +206.3 kJ/mol

CH 4 + CO2 ⇄ 2CO + 2H2, Δh0 = +247.5 kJ/mol

(7)

The implementation of the SOEC model in Aspen Plus is based on the flowsheet introduced by Cinti et al. [19]. A scheme of the modeling approach is depicted in Fig. 4. On the cathode side, the feed stream is first brought into chemical equilibrium in an RGIBBS block, representing the reforming reaction of recycled tail gas. Subsequently, water and carbon dioxide are converted to hydrogen and carbon monoxide in an RSTOIC block. Produced oxygen is separated from the product stream and routed to the anode side from where it is purged by an air stream. The second equilibrium block on the cathode side represents the electrodes catalytic activity. The main reaction occurring is the water-gas shift reaction. At high conversion rates and pressures, also methanation can occur [32]. Three different flowsheets for the implementation of the SOEC were used in this study. In each of the three approaches, counter-current heat exchangers are used to pre-heat the educts on cathode and anode side of the SOEC. The flowsheets differ in the way the thermal energy obtained from by-product combustion is integrated into the syngas production step. The first approach depicted in Fig. 5 is based on the process introduced by König et al. [23] and is comparable to the demonstration plant built by Sunfire [3]. In this concept, water is converted into hydrogen on the cathode side of the SOEC. The obtained hydrogen is converted into syngas together with carbon dioxide and a part of the tail gas in a tube-and-shell-type reverse water-gas shift reactor. The other part of the tail gas is mixed with enriched air and combusted within the shell of the rWGS reactor to supply thermal energy to the reforming and rWGS reactions (in contrast to the Sunfire process, which incorporates an rWGS reactor that is electrically heated). The exhaust gas is then used to pre-heat the educt streams on anode and cathode side of the SOEC. In order to ensure a constant oxygen molar fraction of xO2 = 0.5 at the anode outlet, the Aspen native function ‘design specification’ was used. It allows for the automatic adjustment of the air flow to the cell in order to have the target variable xO2 reach the prescribed value. In the second syngas production scheme (HI1), depicted in Fig. 6a, co-electrolysis and internal reforming are applied. The scheme for by-

(3)

is the main route of carbon dioxide conversion [26,27], others also include the direct conversion of carbon dioxide [28–30]. Even though all of these models were sufficiently accurate to describe experimental data, no consensus on the question of direct carbon dioxide conversion was found. A variance in cell characteristics would explain these fundamentally different approaches showing very reliable results. For the process concepts incorporating a long recycle of by-products into the SOEC, the influence of internal reforming of short-chained hydrocarbons

CH 4 + H2 O⇄ CO + 3H2

ΔḢ . ṅO2 − 2F

(6)

(4) (5)

on cell performance would be important. To the best of the authors’ knowledge, there are no models available, which account for this influence. The internal reforming of by-products was shown to improve overall process efficiency [19]. Additionally, also a reduction of system complexity can be achieved for two reasons. Firstly, no additional reforming reactor is needed and secondly the by-products and unconverted syngas ensure a reducing atmosphere on the cathode side of the cell, which prevents cell degradation. The performance of an SOEC cell is significantly influenced by the gas composition on the electrodes, the current density and the cell geometry [31]. Some of the existing process models excluding internal reforming are based on specific SOEC stack data used in the respective study. This results in either semi-empirical SOEC models that break down the overpotential into ohmic, activation and concentration losses (e.g. [29]) or empirical models. In the latter case, the model is based on curves of the area specific resistance that were fitted to experimental 312

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exhaust gas

(Fig. 6b). It is basically similar to the scheme depicted in Fig. 6a, with the exception, that part of the tail gas is combusted with fresh air (mass flow is adjusted in the simulation to achieve a constant air ratio of λ = 2 in the combustion chamber) instead of enriched anode offgas. The hot exhaust gas of the combustion is fed to the anode side of the SOEC, which allows for the convenient introduction of heat to the electrolysis and reforming reactions. Additionally, this setup allows for pre-heating of anode and cathode feed streams in a single step by using a tube-and-shell-type apparatus. The cathode feed stream is fed through the tubes, while part of the tail gas is combusted within the shell. If a co-current flow is employed both streams exit the pre-heating step with a limited temperature gradient, which reduces thermal stress within the SOEC. Besides the advantageous introduction of heat to the SOEC, the scheme depicted in Fig. 6b leads to a reduction of the oxygen content of the anode feed gas and therefore significantly lowers the electrochemical potential of the system [33]. This allows for a decrease in cell area while maintaining thermoneutral conditions. An additional benefit of this mode of tail gas utilization is the reduced number of heat exchangers, which reduces capital expenditures as well as thermal losses. A ‘design specification’ was used for all syngas production schemes to ensure a constant hydrogen-to-carbon monoxide ratio of H2/CO = 2 for various operating conditions. While the amount of carbon dioxide provided to the process is constant, the fresh water stream that is supplied to the SOEC is varied in order to influence the hydrogen-tocarbon monoxide ratio. Since the syngas composition is determined by a thermodynamic equilibrium either in the SOEC or in the rWGS reactor, a change in the ratio between water and carbon dioxide supplied to the process will result in a shift in the hydrogen-to-carbon monoxide ratio. At the relevant operating conditions of ϑ ≈ 800 °C and p = 1 bar, the hydrogen-to-carbon monoxide ratio is roughly equal to the ratio of water and carbon dioxide supplied to the process. Even though cobalt-based processes for the production of waxes are generally operated at lower hydrogen-to-carbon monoxide ratios [34], a value of H2/CO = 2 was chosen in this work to aid comparability with literature processes [19,20,23]. These concepts were used as a starting point for the process development and a benchmark for the finished process.

air

anode cathode H2O

rWGS reactor CO2

syngas

tail gas Fig. 5. Process scheme for the syngas production via H2O electrolysis and rWGS reaction including thermal by-product utilization according to König et al. [23].

exhaust gas air

anode tail gas

cathode combustion

H2O, CO2 syngas

a)

exhaust gas air tail gas

2.2.2. Hydrocarbon synthesis unit and product separation The Fischer-Tropsch synthesis has gained increasing interest in the last decades as it provides an alternative pathway to hydrocarbons. The necessary syngas can be produced from a variety of sources like coal, natural gas, biomass, biogas or CO2 and H2O [35–38]. Cobalt is the preferred catalyst for the production of high-quality waxes, because it offers high chain growth probabilities and a limited formation of byproducts such as oxygenates and olefins [34]. Even though the product spectrum for cobalt catalysts is significantly narrower than in case of iron [39], the underlying mechanism of cobalt-based Fischer-Tropsch synthesis is very complex and subject to ongoing research [40,41]. Nevertheless, several kinetic approaches for Fischer-Tropsch synthesis were proposed in literature, ranging from microkinetic models [42,43] to simpler power rate law, Langmuir-Hinshelwood and Eley-Rideal type models [44]. However, the effective kinetics as well as the product selectivity are highly influenced by intraparticle diffusion limitations [45,46]. Therefore, the selectivity of the Fischer-Tropsch synthesis not only depends on the chosen catalyst, but also on catalyst shape and on reactor type (e.g. in terms of temperature distribution). For the modeling of a complete process, it is much more convenient to use a simpler approach for the description of the Fischer-Tropsch synthesis step. This is the reason why the well-known Anderson-Schulz-Flory distribution is predominantly applied for the modeling of processes based on FischerTropsch synthesis [19,23,37]. The key parameter of this approach is the chain growth probability α, describing the chain length distribution of Fischer-Tropsch products in good approximation, even though some deviations from the ideal distribution occur in praxis [47,48]. However, the deviations for C1 and C2 components have a rather low influence for

anode cathode

H2O, CO2

b)

combustion

syngas

Fig. 6. Process schemes for syngas production via co-electrolysis including thermal byproduct utilization: (a) HI1 according to Becker et al. [20]; (b) HI2 introduced in this work.

product utilization is derived from the process proposed by Becker et al. [20]. In this case water, carbon dioxide and part of the tail gas are mixed, pre-heated and fed to the cathode side of the SOEC, where the components are electrolyzed and reformed. The other part of the tail gas is combusted with enriched air to pre-heat the educt stream of the cathode side. Combustion and pre-heating are conducted within the same shell-and-tube-type device. This reduces thermal losses and allows for pre-reforming in case of high hydrocarbon contents within the feed gas by adding a catalyst bed to the reactor tubes. Again the air flow is adjusted to ensure a constant oxygen molar fraction of xO2 = 0.5 at the anode outlet. The hot offgas of the combustion step is used to pre-heat the fresh air that is fed to the anode side. This scheme allows for a very efficient combustion, since it is conducted with enriched air. In this paper, a novel heat integration approach (HI2) is introduced 313

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the proposed process, because of their recirculation into the syngas production step (Fig. 3). Another known deviation is the higher chain growth probability for long-chained hydrocarbons C10+. For a practical realization of the proposed process, this would mean that the amount of waxes produced is higher than calculated here. This would be beneficial for economic feasibility, as shown in Section 3.2. Therefore, the application of the Anderson-Schulz-Flory distribution is a viable approach for the modeling study presented here. A reasonable assumption for chain growth probability was reported to be α = 0.9 [49,50]. In conventional multitubular fixed-bed reactors (e.g. applied by Shell [51]), the conversion of syngas is limited to values in the range of XCO = 0.35–0.50 [52–54], mainly due to issues with heat removal from the exothermic reaction. However, for the base case of the simulation the rather high value for conversion of XCO = 0.8 was chosen. This is due to two reasons: Firstly, the authors of this study believe that novel decentralized production processes like the presented one necessitate novel reactor concepts to reach economic feasibility. Results from literature show, that for such concepts very high conversions per pass with a suitable heat removal are achievable [55–57], but also conventional slurry bubble columns offer higher conversions than fixedbed reactors [34,58,59]. A second reason is, that a conversion of XCO = 0.8 was applied in several other process modeling studies [20,60,61]. Therefore, the choice allows for a better comparability to these works. However, XCO was also varied in the present modeling study (Section 3.1.3), showing the influence of the conversion in the Fischer-Tropsch reactor on overall process efficiency. Additional information on the modeling of the synthesis unit and the product separation in Aspen Plus are given in Ref. [37].

Table 2 Standard operating conditions for the process evaluation. Parameter ϑSOEC pSOEC RU ϑcond1 ϑFTR pFTR XCO

thermal loss . heat transferred

ϑcond3 RS RL ΔTHE

Pch,out useable power output . = total power input Pel,in

ṅC ,products molar flow of carbon in the product stream = molar flow of carbon in the educt stream ṅC ,educts

800 °C* 1 bar* 0.75* 15 °C 250 °C 25 bar 0.8* 0.9 250 °C 50 °C* 0* adjusted by ‘design specification’* 75 K 0*

* Value was varied in Section 3.1.3.

all calculations. For ϑws = 15 °C the fraction of condensed water was x H2,cond > 0.99 for all pressure levels. Literature processes assumed a reactant utilization of RU = 0.5–0.9 [19,63,64]. An experimentally validated reactant utilization of RU = 0.75 was chosen for the process evaluation [65,66]. The recycle ratios for short (RS) and long recycle (RL) were defined as

R=

̇ nrecycle ̇ ntotal

.

(11)

The Aspen Plus process model contains several recycles. These as well as ‘design specifications’ cause calculation loops. In order to achieve reliable results ‘design specifications’ were always nested outside, tear streams inside. The tolerance for the outermost calculation loop was set to 10−5. With every further level of nesting, the tolerance was decreased by one order of magnitude. Simulations were carried out with different initial values to exclude the occurrence of local minima.

(8)

3. Results and discussion 3.1. Process efficiency As already stated by Cinti et al. [19], tail gas utilization is of upmost importance in order to achieve a high process efficiency. To study the potential of tail gas recirculation, the recycle ratios of short and long recycle were varied and process efficiency was calculated. At this point, it was assumed that the purge gas was not utilized and the energy demand of the plant (electrolysis as well as feed pre-heating) was met entirely by means of electricity. 3.1.1. By-product recirculation Three different means of tail gas recirculation were investigated. Firstly, a short recycle of tail gas into the Fischer-Tropsch reactor was used. This allows for an increase in overall conversion of syngas within the synthesis step. Secondly, a long recycle of the tail gas into the syngas production step was investigated. This allows for the conversion of by-products back into syngas by reforming reactions on the SOEC cathode. Lastly, both recirculation concepts were combined. As depicted in Fig. 7, tail gas recirculation allows for a significant improvement of efficiency, which is consistent with the findings of Cinti et al. [19]. Due to the high once-through conversion of the FischerTropsch reactor of XCO = 0.8 the amount of unconverted syngas in the tail gas is limited which leads to a slightly lower increase in efficiency if

(9)

Furthermore, the carbon efficiency

ηC =

Minimum temperature difference in heat exchangers Heat loss coefficient

Ψ

All process steps including an exchange of thermal energy are considered to be subject to thermal losses. The hotbox was assumed to be subject to constant heat losses, which were estimated according to Ref. [62]. For high temperature heat exchangers, a ‘hot outlet, cold inlet temperature difference’ specification was chosen to attain comparable results for a wide range of recycle conditions. The specified temperature difference was set to ΔTHE = 75 K. All compressors needed to achieve the desired pressure levels were modeled with an isentropic efficiency of ηisentropic = 0.72. The model comprises the simplification, that pressure drop in the various process steps is not considered. For the evaluation of the process, the energetic efficiency was defined as

ηen =

SOEC temperature SOEC pressure Reactant utilization in the SOEC Temperature of water condensation Fischer-Tropsch reactor temperature Fischer-Tropsch reactor pressure Per pass conversion in the FischerTropsch reactor Chain growth probability Condensation temperature of the wax fraction Condensation temperature of the liquid product fraction Recycle ratio of short recycle Recycle ratio of long recycle

α ϑcond2

2.2.3. Process As shown above, an SOEC-based synthesis process consists of several high temperature process steps that are accompanied by heat losses. The heat loss in heat exchangers was accounted for by the introduction of a heat loss coefficient:

Ψ=

Value

(10)

was calculated to assess the potential of CO2 utilization. An overview over the standard operating conditions of the main process steps is given in Table 2. The temperature of the water separation step was set to ϑcond1 = 15 °C. Technically, such low condensation temperatures are not feasible. However, this value was chosen in order to allow a nearly complete condensation of water for all pressure levels studied and therefore ensure comparable conditions for 314

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0.6

0.60

short recycle long recycle short and long recycle

0.56

0.5

0.4

0.3

xN

2

η en

0.52

0.48 0.2 0.44

0.1

0.40

0.0 0.0

0.2

0.4

0.6

0.8

1.0

R Fig. 7. Energetic efficiency ηen and nitrogen fraction xN2 in the Fischer-Tropsch reactor feed in case of an impure CO2 feed in dependence of the recycle ratios of long and short recycle (Eq. (11)).

Fig. 8. Energetic efficiency ηen for different heat integration concepts in dependence of the heat loss coefficient Ψ for thermoneutral operation of the SOEC.

a short recycle is implemented compared to a long recycle. The highest efficiency can be achieved when both recycles are combined. The efficiency of ηen = 0.537 for an almost closed loop process (RL = 0.95) is consistent with an efficiency of ηen = 0.572 for a closed loop process calculated by Cinti et al. [19]. Though a significant increase in efficiency can be gained by recycling a large share of the tail gas, there are problems associated with high recycle ratios. To illustrate the accumulation of inert components, the simulations series was repeated with a 5% share of inert nitrogen in the carbon dioxide feed stream. The resulting molar fraction of nitrogen in the Fischer-Tropsch reactor feed stream is depicted in Fig. 7. Because of the significant accumulation of inert components, the combination of short and long recycle is not technically feasible for the investigated process concept. While the recycle ratios can be varied independently, for the implementation of a short recycle an additional compressor is needed in order to reach the required pressure level, which increases capital expenditures. Therefore, the short recycle was discarded, because not only nitrogen, but also compounds like carbon dioxide and short-chained hydrocarbons tend to accumulate, which adversely affects the Fischer-Tropsch synthesis and product separation. Additionally, the long recycle allows for the convenient introduction of reductants to the anode side of the cell to avoid cell degradation. Therefore, only the long recycle to the syngas production step was included in the modeling study presented in the next sections. In order to reach a high energetic efficiency as well as a high carbon efficiency, the application of recycles is essential for Power-to-X processes. Therefore, Fig. 7 also highlights that the CO2 streams used must be of high purity in order to allow for the technical realization of a Power-to-X process.

step generally shows the lowest overall process efficiency. This is due to the large amount of water that has to be evaporated to feed the SOEC and the subsequent rWGS step. Additionally, the conversion of carbon dioxide in the WGS step is significantly lower than the reactant utilization of RU = 0.75 that was chosen for the co-electrolysis process schemes. Furthermore, the rWGS concept also shows the steepest gradient with rising thermal loss coefficient Ψ. This is due to the large number of heat exchangers necessary to pre-heat the feed streams. The scheme HI1, which is based on the approach of Becker et al. [20], shows a higher efficiency and a lower gradient. Decisive factors for this are the higher conversion of carbon dioxide in the co-electrolysis step and the lower amount of heat exchangers. Because the heat released by the combustion step still has to be transferred by a heat exchanger to pre-heat the feed stream, accompanied with heat losses, the concept HI1 shows lower efficiency than the concept HI2 proposed in this work. A second mean of heat integration is the utilization of the excess heat generated during the Fischer-Tropsch synthesis. Significant amounts of heat are released at ϑFTR = 250 °C, which can be used to produce pressurized, superheated steam as described by Verdegaal et al. [3]. This way the heat to evaporate the feed water does not have to be supplied by the combustion step, which results in smaller purge stream and subsequently in a larger recycle stream towards the syngas production step. All schemes show a significant increase in efficiency. Due to the largest water demand, the rWGS process shows the most substantial increase in overall process efficiency. However, at a reasonable assumption of Ψ = 0.1 the highest efficiency ηen can be reached with the concept HI2, which shows an increase in efficiency of 7.8 percentage points, while requiring the lowest number of apparatuses. This results in the lowest capital expenditures as well as the highest efficiency. Therefore, the concept HI2 is considered the most suitable for the investigated application and was therefore chosen for further investigation.

3.1.2. Heat integration As already stated, a closed loop process is not technically feasible. Therefore, a certain purge stream has to be discharged from the process. For two reasons this stream cannot be vented into the atmosphere. Firstly, the stream contains poisonous carbon monoxide and therefore is subject to emissions regulations. Secondly, the stream still contains significant amounts of chemical energy. A possibility to meet both requirements is the combustion of the purge gas. The products are harmless carbon dioxide and water. The combustion heat can be utilized to meet the high-temperature heat demand of the process. Three different approaches for tail gas combustion (Section 2.2.1) were modeled and compared with constant boundary conditions. The results for a variation of the thermal loss coefficient Ψ are depicted in Fig. 8. The state-of-the-art process using water electrolysis and an rWGS

3.1.3. Operating points of the main process steps For the process concept HI2, different boundary conditions were varied. For each variation, the other operating conditions were kept at the values given in Table 2. The cell temperature and pressure were found to have the largest impact on process efficiency as depicted in Fig. 9. Due to the coupling of combustion and SOEC, several overlaying trends can be seen. Firstly, when temperature is varied, the amount of heat that is needed to pre-heat the feed streams changes, which subsequently also results in differing heat losses. This leads to a declining 315

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Fig. 10. Energetic efficiency ηen and carbon dioxide content in the syngas xCO2 in dependence of reactant utilization RU for the process concept HI2.

Fig. 9. Energetic efficiency ηen in dependence of SOEC pressure and temperature for the process concept HI2. The black area marks operating conditions where carbon is formed in thermodynamic equilibrium.

downstream synthesis step. For the reasons mentioned above, a high reactant utilization is considered impractical. As values as high as RU = 0.9 are close to operating conditions leading to carbon formation, a lower value of RU = 0.8 was chosen for the calculations in this section. This value has been experimentally validated [21,68] and is close to the value of RU = 0.767 reported by Becker et al. [20]. Another factor influencing process efficiency is the conversion of carbon monoxide and hydrogen in the synthesis step. As depicted in Fig. 11, the conversion has a significant influence on the overall process effciency. As conversion is increased, less unconverted syngas is recycled into the syngas production stage, which significantly decreases the thermal energy demand for educt conditioning and therefore increases efficiency. State-of-the-art Fischer-Tropsch processes based on cobalt are currently operated at rather low concersions in the range of XCO = 0.35–0.50 [52–54]}, which results in a overall process efficiency in the range of ηen = 0.48–0.52. By incorporating novel reactor concepts and improved catalysts, significantly higher conversions have been reported [55–57]. This would allow for an increase in efficiency up to ηen = 0.55. The results in Fig. 11 show that the application of novel reactor concepts is beneficial for Power-to-X processes. Important for technical realization are comparable capital expenditures to stateof-the-art reactors (Section 3.2).

efficiency with rising temperature at low pressure. At high pressure and low temperature, carbon is formed in thermodynamic equilibrium. At temperatures beyond the boundary of carbon formation, a large methane fraction in the syngas hinders efficiency. With a further temperature increase, methane formation is inhibited and efficiency increases. At intermediate pressure, these two mechanisms overlay. There are two different mechanisms depending on pressure as well. Firstly, the electricity needed for syngas pressurization declines rapidly with rising pressure. This is due to the possibility of pressurizing water in liquid state, which saves a significant amount of electricity. Additionally, the reforming of recycled tail gas is a volume increasing reaction. Therefore, the pressurization subsequent to the SOEC has a higher electricity demand. At high pressure levels methanation occurs, which also leads to a decline in efficiency. These overlying effects result in an efficiency maximum at p = 2 bar and ϑ = 800 °C. Even though these boundary conditions allow for the highest efficiency, they are not economically feasible. Pressurized operation of an SOEC requires a pressurized container, an additional compressor and a sophisticated gas dosing to avoid a pressure gradient between anode and cathode side of the cells [67]. This leads to an increase in capital expenditures and is therefore not reasonable. Therefore, for a technical realization the small efficiency penalty when operating at ambient pressure should be acceptable in order to minimize capital expenditures and system complexity. A further important parameter is the reactant utilization RU in the SOEC. Since the cell voltage increases rapidly for high reactant utilizations [68]—which would result in cell degradation—the conversion has to be limited. In process modeling studies the reactant utilization was varied between RU = 0.5–0.9 [19,64] but only values up to RU = 0.81 were experimentally validated [68]. As depicted in Fig. 10, the reactant utilization has a large impact on overall process efficiency. Due to less water being evaporated to supply steam as well as a lower amount of carbon dioxide within the recycle, efficiency increases significantly with increasing reactant utilization. This effect decreases at high reactant utilization. Another advantage of a high reactant utilization is a lower carbon dioxide content in the syngas (Fig. 10), which is favored since high carbon dioxide contents in the syngas facilitates the synthesis of shortchained products [69]. A molar fraction of xCO2 < 0.2 is assumed to not cause reoxidation of the applied cobalt catalyst [70], which is achieved at a reactant utilization of RU = 0.5. This value for RU therefore marks the minimum utilization, where additional syngas upgrading steps are not necessary. At higher values for reactant utilization the carbon dioxide fraction declines, which is beneficial for the

Fig. 11. Energetic efficiency ηen in dependence of conversion XCO in the Fischer-Tropsch reactor for the process concept HI2.

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Table 3 Energy and mass balance of the process depicted in Fig. 3 for the established operating point. Energy balance Apparatus Syngas production SOEC Water condensation Compressor Syngas conditioning Synthesis FTR waste heat Product separation cond1 cond2 Exhaust gas Products Fig. 12. Energetic efficiency ηen and revenue in dependence of the temperature level of the liquid separation step cond3 for the process concept HI2.

Mass balance Value

Unit

53710.0 3018.0 5588.5 3529.5

kWel kWth kWel kWth

2147.3

kWth

0 4449.1 3207.1 −42779.8

kWth kWth kWth kWchem

Stream Inlet Carbon dioxide Fresh water Fresh air Intermediates Dry syngas Recycled water Products Liquids Waxes Exhaust gas

Value

Unit

10.82 5.22 11.08

t/h t/h t/h

14.25 7.19

t/h t/h

−1.67 −1.49 −23.96

t/h t/h t/h

Table 4 Summary of the overall process efficiencies for the considered processes.

Of major importance for process efficiency as well as product composition is the temperature level of the condensation steps. As stated in Section 2.2.3, the wax fraction is separated at reaction temperature by a phase separation at the reactor exit. The liquid fraction is separated in another condensation step. The temperature level of this step can be varied. The impact of the condensation temperature on the overall process efficiency ηen is depicted in Fig. 12. It is obvious that a lower temperature, which favors the condensation of short-chained components, leads to a higher overall process efficiency. The disadvantage of a low condensation temperature is the unfavorable product composition. At low temperature levels, many shortchained components are separated from the product stream. Since the liquid fraction is of lower value than the wax product (Section 3.2), the revenues of the process decline. To maximize profits, the process can be optimized towards the production of waxes as depicted in Fig. 12. To achieve this, part of the liquid fraction could be again converted into syngas and subsequently into waxes. However, reforming of longchained hydrocarbons is challenging as the risk of soot formation increases [71]. Carbon deposits on the catalyst sites of the SOEC cathode would significantly inhibit catalytic activity, which would have an adverse influence on the SOEC performance. In case of recirculation of long-chained hydrocarbons, a pre-reformer prior to the SOEC, as mentioned above, is of vital importance. Since such a process has not been demonstrated so far, it was not included in the process concept. Therefore, the condensation temperature was set to ϑcond3 = 80 °C for the following simulations to avoid long-chained components in the recycle stream.

Reference

ηen

Becker et al. [20] König et al. [23] Cinti et al. [19] This work

0.51 0.446 0.572 0.68 (0.618*)

* including thermal losses with Ψ = 0.1.

comparison. This value is a much better approximation of the efficiency that can be reached by an actual Power-to-X plant. Furthermore, the process concept developed in this work contains fewer heat exchangers compared to the processes introduced by Becker et al. [20] and König et al. [23]. This results in much smaller thermal losses as well as a significant decrease in capital expenditure. Because the exhaust gas of the combustion is vented into the environment after its thermal energy has been used, carbon is removed from the process. At the aforementioned operating point, a carbon efficiency of ηC = 0.86 is reached, which is significantly higher than literature data for the process by König et al. [23] with ηC = 0.73.

3.2. Economic feasibility A study of economic feasibility was conducted assuming an exemplary carbon dioxide emitting process as carbon source. The scale was chosen according to Becker et al. [20] with a mass flow of carbon dioxide of ṁCO2 = 95,000 t/a. The capital expenditures for various process steps were taken from literature data and are presented in Table 5. Since the plant has to be located in close proximity to a carbon source, the indirect costs were estimated for the case of the extension of an existing plant according to Peters et al. [72] with a depreciation period of 20 years. Even though significant progress has been made in SOEC long-term stability [73–75], the degradation of the stacks during operation still has to be considered. Due to this, the capital cost for syngas combustion was split in two parts. The cost for electrolyzer stacks was incorporated into the operational expenditures similarly to the cobalt catalyst used in the synthesis step. In this case, a stack lifetime of ten years was assumed. Only the periphery of the SOEC stacks, which accounts for a third of the cost for the syngas production, is counted towards the capital expenditures. The specific CAPEX of the plant (194.000 €/bpd) are significantly higher than for commercially available GtL processes by Velocys (90.000 €/bpd) [79] and Shell/Pearl (122.000 €/bpd) [80]. Considerably lower values in the range of 25.000–44.000 €/bpd [81,82]

3.1.4. Discussion The boundary conditions of the process were updated to incorporate the results of the one-dimensional optimization conducted in Section 3.1.3. The adjusted boundary conditions are the reactant utilization, which was set to RU = 0.8 and the temperature of the liquid condensation step with ϑcond3 = 80 °C. Additionally, the utilization of the reaction heat of the Fischer-Tropsch synthesis for water evaporation was incorporated in the model for the calculations. All other conditions were kept as stated in Table 2. Using the updated boundary conditions, the process was evaluated. The mass and energy balance is given in Table 3. A feed stream of ṁCO2 = 95,000 t/a was chosen in accordance to [20]. The derived process efficiency as well as literature data for comparison are listed in Table 4. As the literature data does not account for thermal losses, they are compared to the loss-free results (Ψ = 0) calculated in this work. It can be seen that the proposed process shows a significantly higher efficiency in comparison to the literature data. Additionally, an efficiency for a more realistic assumption of Ψ = 0.1 was given for the aid of 317

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Table 5 Summary of capital expenditure data for the main equipment components. Component

Base cost/m€ 2015

Base scale

Scaling factor

Installation factor

Final installed cost/m€ 2015

Reference

SOEC periphery Tail gas combustion Heat exchanger 1

0.62 158.96

1 MWel 65000 m3/h

1 0.8

0.33 1

4.77 11.57 0.47

[76] [77] [78]

0.20

[78]

1.96 3.07 3.24 1.60 15.34 46.02

[61] [60] [60] [61] [72]

Heat exchanger 2 Syngas condensation Compressor Fischer-Tropsch-reactor Product condensation Indirect cost* CAPEX

( 130 ( 130

AHEX 0.78 0.093 AHEX 0.78 0.093

) )

4.09 0.42 18.93 4.09

47.5 t/h 0.4 MWel 2.7 kmol/s 47.5 t/h

0.8 0.67 0.85 0.8

2.5 2.5 2.5 2.5

* indirect costs of a chemical plant amount to 50% of the installed equipment cost according to Peters et al. [72].

40

were reported for the Sasol/Oryx plant. However, these values are difficult to put into context, because it is believed that the plant output has been significantly lower than the reported nominal values for most of the plant lifetime [80]. The higher CAPEX for the process proposed here is due to the smaller scale of the plant. Additionally and more importantly, the SOEC stacks are significantly more expensive than a reformer, which leads to a more costly syngas production. The basis for the estimation of the operational expenditures is based on data of a Velocys GtL plant [37]. The values were adjusted to take into consideration the changed syngas production step and smaller scale. The data is presented in Table 6. As already stated, literature processes focus on the production of fuels. This causes a significant discrepancy between the revenue of 3.3 ct/kWhchem and the production costs of 11.5–48.4 ct/kWhchem depending on electricity price and availability [20,23]. The assumed revenues for waxes (2 €/kg) are much higher than for fuel with 0.45 €/ l, which could result in significant improvements in economic feasibility. Therefore, using the aforementioned data the production costs for co-production of fuels and waxes were calculated (Fig. 13). Costs and availability of electricity were found to be the most important factors in overall production costs. Fig. 13a shows that economic feasibility can be reached if renewable electricity is available at low costs and if waxes are focused as a marketable product fraction. Currently, high availabilities of cheap renewable electricity is limited to a rather small amount of countries. However, the existing plant of Carbon Recycling International [2] in Iceland and the announced plan of Sunfire and Nordic Blue Crude to build a Power-to-X plant with an electric power input of Pel = 20 MW in Norway [83,84] highlight potential countries for a market entry of the technology. In the mid-term future, the technology could also become attractive for countries with considerably lower capital and operational expenditures than assumed in this feasibility study, e.g. India. For a commercially viable production of synthetic fuels in countries like Germany, a significant reduction of CAPEX and OPEX is necessary. Especially the costs for the SOEC stacks have a huge influence on the production costs. Therefore, a reduction of production costs is of vital importance for future applications. In order

production costs / ct/kWh chem

2000 h

OPEX/€/brl

Workforce Catalyst Utilities Maintenance Overheads, other SOEC Electricity OPEX

10 7 1 10 5 31.3* Varied 64.3 + electricity cost

20 revenue waxes

mean revenue

10

revenue gasoline/diesel 0 0

a)

5

10

15

20

electricity costs / ct/kWh el 40

production costs / ct/kWh chem

waxes only various products

30

20 revenue waxes

mean revenue

10

revenue gasoline/diesel 0

b)

Table 6 Summary of the operational expenditures for the proposed process. Component

4000 h 6000 h 8000 h

30

0

5

10

15

20

electricity costs / ct/kWh el

Fig. 13. Production costs in dependence of electricity costs. (a) Influence of availability for the production of various products (ϑcond3 = 80 °C). (b) Influence of product fraction (various products – ϑcond3 = 80 °C; waxes only –ϑcond3 = 250 °C) for an availability of 8000 h.

to allow for a market entry, reduced electricity prices for Power-to-X plants might be a possible governmental tool for the support of this important technology. For a further insight into the influence of product composition on economic feasibility, the production of waxes as the only product was investigated. The recirculation of liquid product fractions hinders

* SOEC lifespan of 10 years was assumed.

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offers the best conditions for a future application due to the abovementioned reasons. The paper delivered valuable specifications (e.g for conversion in SOEC and Fischer-Tropsch synthesis) with regard to a technical realization. The advantages of atmospheric operation of the electrolysis and the necessity for high purity CO2 feed streams were also highlighted. Therefore, the work provides a solid base for further development of this technology towards technical demonstration and industrial application.

overall efficiency, which results in higher production costs. However, since the increase in revenues is higher, the break-even costs for electricity can be raised from 3 to 5.5 ct/kWhel, as depicted in Fig. 13b. However, this would require significant changes to the process concept, foremost a pre-reformer for the conversion of the liquid hydrocarbon fractions to avoid soot formation within the SOEC [71]. The importance of the CO2 source on economic feasibility was already highlighted in Section 1. Because of the great variety of different CO2 sources and due to the great span of published separation costs, the CO2 separation is difficult to include into the economic feasibility study. However, to give an indication of the influence, the increase of production costs for CO2 separation costs of 200 €/tCO2 was calculated. The production costs in this case would increase by 6.6 ct/kWhchem. A comparison with Fig. 13 shows, that this would have a considerable influence on economic feasibility. It is also clear that current CO2 certificate costs in Europe of about 6 €/tCO2 are no strong driver for the application of Power-to-X plants. Because of the significance of this technology for a future energy system, other means of market entry support must be found. A potential pathway might be a compensation for avoided CO2 emissions by Power-to-X plants. The results presented here show, that a compensation of 300 €/tCO2 – as proposed by a German alliance of industrial partners including Audi and Airbus (Power-to-X alliance) in November 2017—might allow for a market entry of efficient processes like the one presented here.

Funding The work was funded within the Fraunhofer Lighthouse Project “Electricity as a Raw Material”. References [1] Bailera M, Lisbona P, Romeo LM, Espatolero S. Power to Gas projects review: Lab, pilot and demo plants for storing renewable energy and CO2. Renew Sustain Energy Rev 2017;69:292–312. [2] Carbon Recycling International; Available from: http://carbonrecycling.is/. [3] Verdegaal WM, Becker S, Olshausen Cv. Power-to-liquids: synthetisches Rohöl aus CO2 Wasser und Sonne. Chemie Ingenieur Technik 2015;87(4):340–6. [4] Schemme S, Samsun RC, Peters R, Stolten D. Power-to-fuel as a key to sustainable transport systems – an analysis of diesel fuels produced from CO2 and renewable electricity. Fuel 2017;205:198–221. [5] Brynolf S, Taljegard M, Grahn M, Hansson J. Electrofuels for the transport sector: a review of production costs. Renew Sustain Energy Rev 2017;81(2):1887–905. [6] Bekker M, Louw NR, Jansen Van Rensburg VJ, Potgieter J. The benefits of FischerTropsch waxes in synthetic petroleum jelly. Int J Cosmet Sci 2013;35(1):99–104. [7] Grand View Research. Paraffin Wax Market Analysis, 2014-2025; 2017. [8] Foit SR, Vinke IC, de Haart Lambertus GJ, Eichel R-A. Power-to-syngas: an enabling technology for the transition of the energy system? Angew Chem Int Ed 2017;56(20):5402–11. [9] Buttler A, Spliethoff H. Current status of water electrolysis for energy storage, grid balancing and sector coupling via power-to-gas and power-to-liquids: A review. Renewable Sustainable Energy Rev 2018;82:2440–54. [10] Yang H, Zhang C, Gao P, Wang H, Li X, Zhong L, et al. A review of the catalytic hydrogenation of carbon dioxide into value-added hydrocarbons. Catal Sci Technol 2017;7(20):4580–98. [11] Saygin D, van den Broek M, Ramírez A, Patel MK, Worrell E. Modelling the future CO2 abatement potentials of energy efficiency and CCS: the case of the Dutch industry. Int J Greenhouse Gas Control 2013;18:23–37. [12] Kuramochi T, Ramírez A, Turkenburg W, Faaij A. Comparative assessment of CO2 capture technologies for carbon-intensive industrial processes. Prog Energy Combust Sci 2012;38(1):87–112. [13] Bhattacharya M, Paramati SR, Ozturk I, Bhattacharya S. The effect of renewable energy consumption on economic growth: evidence from top 38 countries. Appl Energy 2016;162:733–41. [14] Yilmaz C, Turek T. Modeling and simulation of the use of direct reduced iron in a blast furnace to reduce carbon dioxide emissions. J Cleaner Prod 2017;164:1519–30. [15] Arens M, Worrell E, Eichhammer W, Hasanbeigi A, Zhang Q. Pathways to a lowcarbon iron and steel industry in the medium-term – the case of Germany. J Clean Prod 2017;163:84–98. [16] Climeworks; Available from: http://www.climeworks.com/. [17] Goeppert A, Czaun M, Surya Prakash GK, Olah GA. Air as the renewable carbon source of the future: an overview of CO2 capture from the atmosphere. Energy Environ Sci 2012;5(7):7833–53. [18] Broehm M, Strefler J, Bauer N. Techno-economic review of direct air capture systems for large scale mitigation of atmospheric CO2. http://dx.doi.org/10.2139/ ssrn.2665702; 2015. [19] Cinti G, Baldinelli A, Di Michele A, Desideri U. Integration of solid oxide electrolyzer and Fischer-Tropsch: a sustainable pathway for synthetic fuel. Appl Energy 2016;162:308–20. [20] Becker WL, Braun RJ, Penev M, Melaina M. Production of Fischer-Tropsch liquid fuels from high temperature solid oxide co-electrolysis units. Energy 2012;47(1):99–115. [21] Cai Q, Luna-Ortiz E, Adjiman CS, Brandon NP. The effects of operating conditions on the performance of a solid oxide steam electrolyser: a model-based study. Fuel Cells 2010;10(6):1114–28. [22] Alenazey F, Alyousef Y, Almisned O, Almutairi G, Ghouse M, Montinaro D, et al. Production of synthesis gas (H2 and CO) by high-temperature Co-electrolysis of H2O and CO2. Int J Hydrogen Energy 2015;40(32):10274–80. [23] König DH, Freiberg M, Dietrich R-U, Wörner A. Techno-economic study of the storage of fluctuating renewable energy in liquid hydrocarbons. Fuel 2015;159:289–97. [24] Tremel A, Wasserscheid P, Baldauf M, Hammer T. Techno-economic analysis for the synthesis of liquid and gaseous fuels based on hydrogen production via electrolysis. Int J Hydrogen Energy 2015;40(35):11457–64. [25] Isaksson J, Åsblad A, Berntsson T. Pretreatment methods for gasification of biomass and Fischer-Tropsch crude production integrated with a pulp and paper mill. Clean

4. Conclusions In this paper, different methods of heat integration and by-product recirculation proposed in the literature were compared with a novel integration concept in a unified simulation environment. The results showed that co-electrolysis allows for the production of syngas with a suitable hydrogen-to-carbon monoxide ratio of H2/CO = 2 in a one-step process while allowing for easy heat integration. An important factor for reaching a high overall process efficiency is the utilization of tail gas. It was found that by combining a long and a short recycle the highest efficiency can be reached. However, this process scheme is not feasible due to the accumulation of inert compounds. Therefore, only a long recycle was implemented because the efficiency gains were larger than in case of a short recycle. The utilization of the purge gas of the process for pre-heating the feed streams of the SOEC was considered. Three different approaches for the integration of combustion heat were modeled—two schemes derived from literature data and one proposed in this work—and compared using a unified simulation environment. Considering thermal losses, the proposed scheme for syngas production showed a significantly higher process efficiency, while necessitating fewer heat exchangers. For said process scheme, a one-dimensional optimization was conducted to detect optimal operating conditions. Incorporating several means of heat integration an overall efficiency of ηen = 0.68 was obtained. Considering heat losses within the process a still very high efficiency of up to ηen = 0.62 can be reached with the process. In summary, the proposed process offers the following advantages in comparison to the state of the art:

• Highest overall and carbon efficiency due to advanced heat integration and by-product recirculation. • Reduction of system complexity and number of necessary reactors and heat exchangers. • Consideration of heat losses. This features allowed for a valuable evaluation of the economic feasibility of Power-to-X processes at the current state of development. It was also highlighted that a market entry of Power-to-X processes in the short and mid-term future is only likely if governmental support is provided. However, from a technical point of view, the proposed process 319

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