Production of synthetic natural gas from industrial carbon dioxide

Applied Energy 260 (2020) 114249

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Production of synthetic natural gas from industrial carbon dioxide a

b

c

b

a,⁎

Remi Chauvy , Lionel Dubois , Paul Lybaert , Diane Thomas , Guy De Weireld a b c

T

Thermodynamics and Mathematical Physics Unit, University of Mons, 20 Place du Parc, 7000 Mons, Belgium Chemical and Biochemical Process Engineering Unit, University of Mons, 20 Place du Parc, 7000 Mons, Belgium Thermal Engineering and Combustion Unit, University of Mons, 20 Place du Parc, 7000 Mons, Belgium

HIGHLIGHTS

has a strong potential in the transition to a renewable network. • Power-to-gas CO capture process allows to reduce the energy consumption. • Advanced integration improves the competitiveness of the global Power-to-Gas chain. • Heat • Attractiveness of Power-to-gas is economically sensitive to electricity cost. 2

ARTICLE INFO

ABSTRACT

Keywords: Carbon capture and utilization Power-to-gas Process simulation Techno-economic evaluation Heat integration

The Power-to-Gas strategy has become a mainstream topic for decarbonization and development of renewables and flexibility in energy systems. One of the key arguments for decarbonizing the gas network is to take advantage of existing network infrastructure, gradually transitioning to lower fossil carbon sources of methane from Power-to-Gas. This work proposes the techno-economic investigation of an integrated system considering an advanced CO2 capture process, in terms of solvent and process configuration, to treat about 10% of a cement plant’s flue gas and convert the captured CO2 into synthetic natural gas using renewable hydrogen generated from a large-scale wind powered electrolyzer. An optimized heat recovery system is proposed, drastically decreasing the external hot utility demand of the CO2 capture unit. In addition, it leads to the production of complementary electricity (about 1.06 MW), reducing thus also the electrical demand of the integrated process. The synthetic natural gas produced has a composition (CH4 92.9 mol.%, CO2 3.7 mol.%, and H2 3.4 mol.%) and a Wobbe index (46.72 MJ/m3), corresponding to specification for gas grid injection at 50 bar in Germany. With an overall system efficiency of 72.6%, the process produces 0.40 ton synthetic natural gas per ton of captured CO2. The cost of the synthetic natural gas produced is higher when compared to the present natural gas market price, but cost reductions and possible commercial use of coproducts like oxygen, represent a likely alternative. Costs are mainly driven by high capital investments (the electrolyzer), and the price of renewable electricity, which is expected to decrease in the coming years.

1. Introduction Following the 2016 Paris Agreement (COP 21) to keep warming below 2 °C above pre-industrial levels, pursuing efforts to limit the temperature increase to 1.5° C, above pre- industrial levels, the European Union (EU) reaffirmed its goal to reduce greenhouse gas emissions by 80–95% compared to 1990 levels by 2050. In order to achieve these goals, large-scale implementation of low-carbon technologies, such as renewable energy sources (RES), and carbon capture utilization and storage (CCUS), are necessary. Numerous EU countries

⁎

are therefore considering these two families of technologies, such as Norway, Germany, and Denmark. Zappa et al. recently conducted a study investigating whether a 100% renewable European power system would be feasible by 2050 [1]. They concluded that a 100% renewable power system would still require significant flexible zero-carbon storage capacity to balance variable renewable energy generation, and CCUS may still be required to achieved the EU’s ambitious climate objective. Nowadays, global wind power capacity is growing continuously (310 TWh per year in 2015 in European countries with an installed

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

https://doi.org/10.1016/j.apenergy.2019.114249 Received 29 July 2019; Received in revised form 21 November 2019; Accepted 23 November 2019 0306-2619/ © 2019 Elsevier Ltd. All rights reserved.

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capacity of 136 GW [1]) as the cost of producing renewable electricity is sometimes even lower than the cost obtained from fossil fuels. Penetration of wind power in the electricity market is already high in several EU countries such as Denmark (51%), and Ireland (24%). Germany also denotes a high share of solar-based electricity (6%) [2]. However, the penetration of renewable sources into the energy system is often limited by the energy storage capability, mainly due to high costs. Therefore, for continuous availability and reliability, development of other efficient grid storage alternatives is needed. In this context, the Power-to-Gas (PtG) strategy is of growing interest for decarbonization and increasing development of renewables and flexibility in energy systems [3]. In PtG applications, when abundant renewable energy is available, mainly (but not necessarily exclusively) excess electricity can be used to produce a storable gas, such as hydrogen produced by water electrolysis. Currently, only 4% of hydrogen is produced by water electrolysis (overwhelmingly by steam reforming), where the total hydrogen production around the world is about 500 bill·Nm3/year [4,5]. This renewable hydrogen is either used directly in industrial applications or in fuel cells [6]. Direct injection of hydrogen into the natural gas grid (up to 10%) or conversion into synthetic natural gas (SNG) by methanation and subsequent injection into the gas grid are widely considered in the literature [2,7,8,9]. In addition to maintain the high heating value of the gas, methanation denotes several advantages, as it avoids adaptation in the gas grid equipment necessary for higher hydrogen contents. Furthermore, compared to pure hydrogen, SNG (87–97 wt% methane) has fewer barriers to implementation, as it is safer, easier to transport and store, and more suitable for industrial applications [10]. It overcomes the facility and energy density issues related to the use of hydrogen, while providing other services [11]. It has to be noticed that the conversion is limited in the term ‘Power-to-Gas’ to the production of hydrogen and methane, the later in the form of SNG. Other converting forms of electrical energy to liquid energy carriers (hydrocarbons such as methanol, dimethyl ether, and Fischer-Tropsch products) are more likely to associate the term ‘Power-to-Fuel’ or ‘Power-to-Liquid’ [12]. Even though hydrocarbons denote a higher volumetric density in comparison to gases, PtG allows the use of the existing natural gas infrastructure as transport and storage medium, offering a market availability of all the system-relevant components. Furthermore, wind-to-SNG offers an energy storage medium for the intermittent wind resource, where the electricity generated by the renewable energy is merged into the grid providing a more constant renewable energy supply, avoiding oversupplying power grids with high production from renewable energy. Concurrently, many industrial processes look for possibilities to decrease their CO2 emissions, where CCUS is considered. Accounting for 2–2.5 GtCO2 per year (5–7% of the total anthropogenic CO2 emissions) [13,14], the cement sector is investigating ways to cut its emissions through different levers such as modern dry-process technology, clinker substitution, free-carbon alternative fuels and CCUS. It is estimated that only about one third of the CO2 released is due to combustion, and this makes the cement industry a specific case compared with other combustion industries. Clinker, main constituent of the cement, is made by heating a homogeneous mixture of raw materials at minimum temperature of 1450 °C. Two thirds of these released emissions, i.e. 550 kgCO2 per ton of clinker, come from the limestone calcination during the decarbonation step in the clinker burning process, and are then necessarily emitted [15,16]. Hence, the capture of CO2 from cement plant’s flue gas and its conversion are key issues. This CO2 can be a source for methanation. Several technologies at different levels of maturity and performances can be envisaged for capturing the CO2 from cement kilns (e.g. oxy-fuel combustion, chilled ammonia technology, adsorptive processes, calcium looping, etc.) [17–20]. The most mature technology is the absorption-regeneration amine-based process, which reached a TRL (Technology Readiness Level) of 9 compared to the lower TRLs of all other technologies. More precisely, a recent study from the IEA

Greenhouse Gas R&D Programme [21] pointed out that among the different liquid absorbents that can be used for capturing CO2 (e.g. aqueous amine (TRL 6–9), amino acid or mixed salts (TRL 6), ionic liquids (TRL 4), water-lean absorbents (TRL 5), precipitating and demixing solvents (TRL 4–6), etc.), aqueous solutions composed of amine blends (e.g. piperazine (PZ) mixed with another amine such 2-amino-2methyl-1-propanol (AMP) or methyldiethanolamine (MDEA)) can be considered as the new reference solvents in place of monoethanolamine (MEA), long time considered as benchmark for such CO2 capture processes [21]. These amines blends and especially activated solutions allow to reduce significantly the energy demand of the system. Regarding the other post-combustion CO2 capture technologies, membrane systems (TRL 5–6) and solid sorbent processes (TRL 6) (e.g. Vacuum Pressure-Swing Adsorption (VPSA), calcium looping, etc.) have also received a growing interest [22–24]. The widespread deployment of pilot and demonstration plants, including both large-scale demonstration projects and commercial size amines-based post-combustion worldwide installations, such as Boundary Dam (Canada), Petra Nova (USA), and Shengli (China), were reviewed by Idem et al. [25]. In the case of the post-combustion CO2 capture by absorption–regeneration processes applied in a cement plant, the CO2 is separated from the exhaust gases of the system by adding a unit to the tail-end of the clinker process. While this end-of- pipe mature technology achieves high absorption efficiency (usually 90%), the main disadvantage is the high energy demand for the scrubbing liquid regeneration (e.g. from 3.2 GJ/t CO2 to around 3.8 GJ/t CO2 with MEA 30 wt% as solvent [26]). Nevertheless, the development of new solvents (new solutions or blends) [21,27], and new process configurations [28,29] are allowing the global performances of the absorption-regeneration technology to reduce its energy consumption to be improved, so that this technology remains attractive and competitive. The Power-to-Gas process chain was first proposed in Japan in the late 1980 s. Recently, a growing interest dealing with this technology has begun, especially in Europe. In a detailed review, Rönsch et al. give an overview of methanation technology and research, focusing on projects with methane as a product [30]. Chauvy et al. maintain that PtG is one of the best alternative processes for CO2 utilization among examined cases in the sensitivity analysis [31]. The interest is mainly driven by the increasing share of renewable energy (wind and solar power), where PtG may provide large-scale and long-term energy storage [32]. Other field of applications include services to balance the loads in electricity networks, a substantial source of fuel for heating and transportation, and a significant contribution to emission reduction targets [33]. Pilot plants are under construction, or even in operation, in several countries, including Germany, Switzerland, Denmark, France, and Japan. In particular, the European Power-to-Gas Platform contains a database of past, current and planned PtG projects in Europe [34]. Numerous reviews of PtG projects present lab, pilot and demo plants for storing renewable energy and CO2 [7,35,36]. Quarton and Samsatli is particularly recommended for more detail on real-life PtG projects, economic assessments and systems modeling [7]. Bailera et al. detail worldwide existing projects and present basic information together with qualitative descriptions of the plants, technical data, budget and project partners [35]. Wulf et al. reference about 130 demonstration projects in operation or planning in Europe (63 projects were in operation by end of May 2018), about 40% of which were developed in Germany [36]. One example project that illustrates the concept of PtG is the e-Gas Project (Werlte, Germany) [8,35]. It demonstrates how renewable energy could be efficiently stored in the existing natural gas network. It consists of methanation combined to an organic waste biogas plant, which produces about 1 kton methane-rich gas from concentrated CO2 and renewable H2. Three alkaline electrolyzer stacks of 2 MW supplied from a 14.4 MW capacity offshore wind park generate the H2. The annual SNG injection to the grid from this process is approximately 3 million Nm3. 2

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Other CCUS route(s) Raw materials Fuel

Cement kiln

A

CO2-separation unit Amine scrubbing

Flue gas

Cement

Scrubbing liquid regeneration Heat

CO2 Excess RE H2O

Electrolysis

Depleted flue gas

H2

B

Excess heat for other application

Methanation CO2/H2

O2

Raw SNG

Gas upgrading

H2O SNG

Grid injection

Fig. 1. Conceptual Flowsheet of the interconnected PtG system comprising (A) CO2 capture unit; (B) CO2 methanation unit.

In the meantime, many studies focus on thermodynamic considerations, the development of new catalysts, the characterization of the reaction kinetics and mechanisms of CO2 methanation [37,38]. For instance, Su et al. proposed a review on the advances in catalytic CO2 hydrogenation to methane, discussing the structure of the catalysts, the preparation methods, the pre-treatment conditions and their components in detail [37]. In addition, Eveloy and Gebreegziabher recently reviewed projected PtG deployment scenarios at regional and distributed scales [39]. They highlighted that even though many types of PtG energy/material integrations are possible, few of them have been incorporated into PtG deployment modelling works. Furthermore, they identified that heat recycling from catalytic methanation reaction products, feed compression trains, and condensed water integrated with CO2 capture from cement production and/or water electrolysis requirements have received limited attention. This lack of experience with the whole PtG system was also highlighted by Quarton and Samsatli [7], Ghaib and Ben-Fares [8], and Thema et al. [40]. It has to be pointed out that the development of PtG technology is also subject to fundamental energy and climate policy decisions. Therefore, only a few studies attempt to evaluate large-scale integrated PtG process performances, showing a large potential for improvement, which helps the decision makers to justify their choices. Becker et al. recently proposed a detailed SNG plant study case where the synergies between large-scale reactors, thermal management and techno-economic considerations were provided [41]. Morosanu et al. proposed a concept based on water electrolysis to produce hydrogen, CO2 capture from ambient air using solid adsorption materials, catalytic CO2 methanation, gas separation, and a single mixed refrigerant (SMR) methane liquefaction process, where mass and energy balances at demonstration scale were provided [42]. To this extent, in-depth modelling and economic analysis of PtG systems are addressed in this paper. The integration of SNG production with an advanced CO2 capture process, mainly in terms of solvent and process configuration, and considering an optimized heat recovery leading to a complementary electricity production, are especially innovative and correspond to the purpose of the present study. The suggested PtG technology is flexible, easily up/down-scalable and modular in order to allow an adjustment to any specific boundary conditions of a distinct application, such as end-product, source of CO2, use of potential by-products, and main purpose of the system (utilization of excess renewable energy, storage, etc.).

part of CO2 coming from a cement plant into SNG, and comprises four implementation steps: (i) an operational cement plant equipped with the Best Available Technique (BAT), where an advanced CO2 separating unit for scrubbing CO2 contained in the flue gas was considered; (ii) an electrolysis unit for splitting water by electrolysis to produce renewable H2 and O2; (iii) a methanation unit for converting the CO2 and the renewable H2 by a methanation reaction to a methane rich gas (SNG); and (iv) a SNG upgrading unit to fulfil the requirements for further injection into the gas grid. Finally, the possibility of heat integration through systematic process-to-process analysis is investigated to reach high energy efficiency and minimize utility costs. It is worth noting that the produced SNG can also be converted into compressed renewable natural gas or liquefied renewable natural gas to serve as a transport fuel. The size of the PtG process is based on a 90 MW large-scale windbased electrolyzer, generating 40 ton of renewable H2 per day. The flue gas to be treated comes from a conventional BAT cement plant producing 3000 tons of clinker per day. It corresponds to a total flow rate of 250,000 m3/h, at 1.20 bar and 40 °C after conditioning (after desulfurization, denitrification, dedusting and cooling), with a CO2 content of 20 mol.%, generating 2475 ton of CO2 per day [43]. Therefore, 9.8% of the cement plant flue gas is treated where 90% of the CO2 is captured, representing a CO2 production of about 219 ton per day. The CO2 capture unit can be upscaled in order to treat the entire cement plant’s flue gas, where the captured CO2 can be geologically stored or converted into other products. Fig. 1 illustrates the conceptual flowsheet of this interconnected PtG system. 2.1. Process modeling and design 2.1.1. Thermodynamics models and simulation tools The units investigated in this study were implemented in AspenTech’s software. Concerning the CO2 capture unit, it was simulated using the Acid Gas Property Package in Aspen Hysys™ v10.0 software. More precisely, the thermodynamic models implemented in such package are the Electrolyte Non-Random Two-Liquid (eNRTL) activity coefficient model for electrolyte thermodynamics in the liquid phase [44] and the Peng-Robinson equation of state for the vapor phase [45]. The Acid Gas Package also includes the physicochemical properties of acid gases (CO2 and H2S), water, amines alone (e.g. MEA, DEA, MDEA, PZ, etc.) as well as several mixtures (e.g. MDEA + PZ, DEA + PZ, etc.). The CO2 capture modelling is based on a rate-based calculation model and comprises a makeup unit operation to automatically compensate water and amine losses, which was validated by Laribi et al. [46].

2. Materials and methods The present PtG integrated process plant was developed to convert a 3

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Aspen Plus™ v10.0 was used for both the CO2 methanation and gas upgrading units considering the Peng Robinson equation of state for the calculation of gaseous properties. Estimates were considered to model the block unit for hydrogen generation.

regenerated at higher temperature for optimal performances [29], a regeneration pressure (Pbottom of the stripper) of 5 bar was considered, leading to a regeneration temperature of 140.8 °C, in contrast to a temperature Tregen of around 120 °C with a conventional MEA 30 wt% solvent. The RVC configuration (applied to the rich solution, i.e. after the absorption step) uses a heat pump effect. As described by Le Moullec et al., the purpose of the RVC process modification is to increase the heat quality provided to the system by enabling the valorization of heat available at a lower quality level or when increasing the quality level is energetically interesting [28]. Thanks to the flashing of the solvent, a gaseous stream mainly composed of CO2 and H2O is produced. This stream is compressed and sent back to the regeneration column in order to reduce the steam demand at the reboiler. At the same time, with the vapor coming from the compressor following the flash tank being quite hot (which could lead to hot spot and degradation problems in the stripper’s bottom if it is reinjected at such temperature level), a second internal heat exchanger is added in order to cool this vapor down to the same temperature level at the bottom of the regeneration column (Tregen ). Such an operation also has the advantage of improving the preheating of the rich solution before entering the stripper. Concerning Intercooling, this is a well-described technique which has already been applied in several pilot or industrial units. It can also be considered in the “absorption enhancement” process modification category. The principle of intercooling is to withdraw the solvent (partially or totally) flowing in the absorber, at a stage n , to cool it down and to send it back to the absorber at the stage n 1. Such a method enables a shift in the thermodynamic gas–liquid equilibrium, and consequently increases the rich loading at the absorber bottom (at a same liquid flow rate) [28]. However, a lower temperature in the absorber leads to reduced chemical kinetics and diffusivities which is mostly compensated by the increase of CO2 solubility leading globally to very little change to the overall mass transfer coefficient. Globally, such a technical option is expected to decrease the reboiler duty, as highlighted by Le Moullec et al. [28]. Regarding the water-wash sections, as described in the IEAGHG report prepared by CSIRO (Australia) [52], scrubbers are used for extracting condensable or soluble vapor from gases. Amine vapors and their degradation products are conventionally captured by water wash or a scrubbing stage, or several stages with demineralized water, acidic water or with special reagents. In the present work, and in accordance with the CASTOR/CESAR pilot unit considered as a reference, two water-wash sections were added to the simulation flow sheet in comparison with the previous model described by Dubois and Thomas [29]. As such, one water-wash section was used at the top of the absorber, consisting of a packed column composed of three-stages of 1 m height per stage, and another wash section was used directly inside the top of the stripper consisting of three stages of packing (same dimensional parameters as for the main column section) below the condenser and above the rich solution inlet, in which the vapor coming from the regeneration meets the liquid coming from the condensation. In the absorber water-wash section, a freshwater make-up is added and also serves as water make-up for the global flow sheet. Only a small complement is added into the make-up unit. It is assumed that the pre-treated flue gas (after de-SOx, de-NOx, dedusting and cooling) entering the CO2 capture process (composition in Table 2) consists primarily of CO2, H2O, N2 and O2. This composition is based on average values coming from the Brevik Cement plant (Norcem company) in Norway. It is important to note that the amount of CO2 in the gas to be treated was maintained at 20.4% for all the simulated cases. Nevertheless, the influence of the inlet gas CO2 content on the absorption–regeneration performances has been discussed in previous works [29,53]. A summary of the simulation results of the implemented process highlighting its energetical benefit in comparison with the conventional process using MEA 30 wt% as solvent, is provided in the Supporting

2.1.2. Renewable H2 production unit Water electrolysis is mainly characterized by the source of electricity, i.e. high temperature steam, solar or wind-based electrolysis, and the solution used as electrolyte, i.e. alkaline (AE), proton exchange membrane (PEM) or solid oxide (SOEC) electrolyzers. An essential block for PtG, electrolyzers need to fulfil some special requirements, including dynamic behavior to follow fluctuating power inputs, high efficiency to avoid energy losses, elevated pressure to reduce compressor costs, as well as long lifetimes. Even though alkaline electrolyzer technology is fully mature, PEM water wind-based electrolysis, which is currently close to commercial deployment, was chosen in this work as it is more compact and best suited to the dynamic load balancing of electricity grids needed with use of intermittent renewable energy. Traditional alkaline electrolyzers, on the other hand, are typically used for continuous steady-state industrial production of hydrogen [47,48]. The comparison of electrolyzer technologies is sum up in the Supporting Information (see SI.2). PEM electrolyzer produces hydrogen at 30 bar with a high purity level of 99.9 vol%, whilst the operating temperature is typically low (50–100 °C) [49]. The specific electrical power of these electrolyzers is estimated at 5 kWh per Nm3, approximately 55.6 kWh per kg H2. A production cost of 4–5 € per kg H2, and a total system cost, including power supply and installation costs, of 1200 € per kW with an operation lifetime of 20 years are considered [41,49]. The oxygen stream from the water electrolysis is produced in high purity, and may be further used in the calcination process to increase the CO2 partial pressure which facilitates CO2 separation (partial oxy-combustion) [17,19,46]. This high purity byproduct can also be traded for many high-tech applications. The electrolyzers are assumed to be co-located with the sources of renewable electricity and CO2, neglecting de facto the efforts to integrate the different units, such as infrastructures, and avoid transport needs. 2.1.3. CO2 capture unit The implemented CO2 capture process was based on the works of Dubois and Thomas [29,50] considering a Rich Vapor Compression (RVC) process combined with an Inter-Cooled Absorber (ICA) and two Water-Wash (WW) sections, and the use of a MDEA 10 wt% + PZ 30 wt % blend as the solvent. The selection of the configuration and the solvent, specifically the use of PZ as absorption activator in comparison with other ones (e.g. (piperazinyl-1)-2-ethylamine – PZEA), based on reaction mechanisms differences leading to different absorption-regeneration performances, are derived from previous simulation results (see Supporting Information for more details). Fig. 2 presents the implemented flow sheet. The installation was designed taking the pilot unit used in the CASTOR/CESAR European Projects as Ref. [51]. The CASTOR/CESAR unit was selected as a reference since all the design and operating parameters in relation with the installation are available in literature [51]. In the current study, this unit was upscaled (columns diameter) in order to treat 24,660 m3/h of flue gas with a CO2 absorption ratio of 90 mol.%, corresponding to a production of 219 ton per day of CO2 (98 mol.% purity). The upscaling was performed to keep the same gas and liquid velocities, 1.29 m/s and 6.70 10−3 m/s, respectively, such as the optimum value for the (L/G) vol ratio, namely around 5.19 10−3 m3/ m3. The dimensions of the absorber and the stripper, and the operating conditions for each column are provided in Table 1. Note that Tabs corresponds to the absorption temperature, at the top of the absorption column, while Tregen refers to the regeneration temperature, at the bottom of the stripper. PZ-based solvents being conventionally 4

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Fig. 2. Aspen Hysys™ flow sheet of the RVC + ICA + WW configuration with MDEA + PZ as solvent.

High pressure and low temperature enhance the CO2 methanation process performance, as denoted in Fig. 3. Below 600 °C, only species involved in the Sabatier reaction (see Eq.(1)) are present, with CO appearing only for higher temperatures (see SI.3.1). Gao et al. underline that carbon deposition is not observed under these simulation conditions for a H2/CO2 stoichiometric ratio of 4, as the water produced during the methanation process reacts with solid carbon [55]. In this work, operating conditions are chosen in such a way that CO and solid C are not produced. In addition, this would allow the simplification of the following gas upgrading step before injection into the grid [56]. The main detailed kinetic law was developed in 1989 by Xu and Froment on the Ni/MgAl2O4 (Ni 15 wt%) commercial catalyst, defined for temperatures between 300 and 400 °C and pressures between 3 and 10 bar without dilution gas that is closer to the industrial implementation of CO2 methanation [57]. A lower temperature (below about 200 °C) may cause the formation of nickel carbonyl, while a high temperature above 500–650 °C may cause carbon deposition (coking), which will deactivate the catalyst [58]. Modelling catalyst reactions in a series of catalytic reactors has been considered, as it avoids the coking and sintering of the catalyst. At least two adiabatic reactors have to be connected in series for a good control of the reaction temperature [59]; the usual approach relying typically on a series of two to five reactors [60]. In this work, the methanation unit consists of four adiabatic multitube fixed bed reactors, featuring a multi-stage cooling system and gas recirculation, allowing the outlet specifications to be achieved while minimizing costs. The Gas Hourly Space Velocity (GHSV) was maintained close to 4000 h−1. The corresponding rate equations are therefore given by Eqs. (4)–(7), where the kinetic parameters and their implementation are discussed in the Supporting Information (see SI.3.2).

Table 1 Main process data equipment for the CO2 capture unit. Parameter

Absorber

Stripper

Diameter (m) Packing height (m) Packing type

2.6 17 (17 × 1 m)

2.6 10 (10 × 1 m)

Random packing IMTP Norton Metal 50 mm MELLAPACK, SULZER, Standard, 125× (for the water wash section) 40 – – 140.8

Tabs (°C) Tregen (°C)

Pbottom (bar)

1.2

5

Table 2 Composition of the conditioned gas to be treated (G = 24,660 m3/h, 1.20 bar, 40 °C). Component

mol frac.

N2 CO2 O2 H2O CO SO2 NO NO2

64.7% 20.4% 8.6% 6.2% 1330 ppm 111 ppm 474 ppm 2 ppm

Information, where the reactions that are included in the Acid Gas Package for PZ and MDEA solvents reacting with CO2 are also presented (see SI.1). 2.1.4. CO2 methanation unit Thermodynamics and kinetics. Methane production by catalytic CO2 hydrogenation (Sabatier reaction) has been used since the beginning of the 20th century. The reaction scheme is given by [54]:

CO2 + 4H2

CH 4 + 2H2 O

CO2 + H2

CO + H2 O

CO + 3H2

CH 4 + H2 O

H0298 = 0 H298

165 kJ/mol

= + 41 kJ/mol

H0298 =

206 kJ/mol

pCH4 × pH22 O k1 r1 = 3.5 × DEN 2 pH2

(1) (2)

r2 =

(3) 5

pCO × p H2 O k2 × p H2 DEN 2

4 ×p pH CO 2 2 K1

(4)

p H2 × pCO 2 K2

(5)

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1.0

100%

0.8

80%

0.6

60%

0.4

40%

0.2

20%

CH4 yield (%)

Molar composition

R. Chauvy, et al.

H₂ CO₂ H₂O CH₄ CO CH₄ yield (10 bar) CH₄ yield (1 bar)

0.0 200

300

400 500 600 Temperature (°C)

700

800

0%

Fig. 3. Molar concentrations of CO2, H2, H2O, CH4 and CO at stoichiometric ratio H2/CO2 = 4, P = 10 bar, and CH4 yield at P = 1 and 10 bar).

pCH4 × p H2 O k3 r3 = 2.5 × DEN 2 pH2

electrolysis, considering the stoichiometric ratio H2/CO2 of 4. A singlestage CO2 compression step is considered technically appropriate to reach the working pressure of 10 bar, while hydrogen is brought to the same pressure. They are then mixed with the recycle stream derived from the upgrading unit and pre-heated to 350 °C before feeding the first adiabatic reactor (REA-1) where the methanation reactions occur. A fraction of the effluent, set at 0.7, is recycled to the inlet of REA-1 to lower the temperature below 650 °C in order to avoid catalyst sintering [61]. Heat removal is ensured to counteract thermodynamic limitations of the reaction as well as to prevent the catalyst bed from sintering the methanation reactors. The reaction mixture is therefore cooled down to 350 °C after each reactor to obtain high CO2 conversion. The heat of the methanation reactions can be partly used to produce steam to regenerate the scrubbing liquid in the CO2 capture unit. After methanation, water vapor has to be separated from the product gas. The outlet of the last reactor (REA-4) is therefore cooled to a temperature of 25 °C to condense the water. A purity of 99.99 vol% is then achieved in the H2O stream. Table 3 presents the main process data equipment for the CO2 methanation unit, while Table 4 introduces the specifications for the four

3 ×p pH CO 2

K3

(6)

where

DEN = 1 + K CO pCO + K H2 p H2 + K CH4 pCH4 +

K H2 O p H2 O p H2

(7)

and with ri referring to the rate of reactions i (reactions (1)–(3), respectively), k1, k3, the rate coefficients of reactions (1) and (3) (kmol bar1/2/kgcat h), respectively, k2 the rate coefficient of reaction (2) (kmol/bar kgcat h), K1, K3 the equilibrium constant of reactions (1) and (3) (bar2), respectively, K2 the equilibrium constant of reaction (2), K CH4 , K CO , and K H2 the adsorption constants of reaction for CH4, CO and H2 (bar−1), respectively, K H2 O the desorption constant of H2O, and p the partial pressure (bar). Process modelling. Fig. 4 shows the process flow sheet of the CO2 methanation unit including the upgrading section. The CO2 coming from the CO2 capture unit is fed at 5 bar with hydrogen coming at 30 bar from the PEM water wind-based

Fig. 4. Aspen Plus™ flow sheet of the CO2 methanation unit including the upgrading section. 6

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Table 3 Main process data equipment for the CO2 methanation unit. Parameter

Table 5 Gas specification for gas grid injection in Germany [56,69,70]. Specification

Reactor operating Pressure (bar) Pressure Loss (%) Reactor operating Inlet Temperature (°C) Catalyst Catalyst density (kg/m3) Bed density (kg/m3) Bed porosity

Parameter

10 2 350 Ni/MgAl2O4 2350 790 0.44

Wobbe Index (MJ/m ) CO2 H2 Dust water

3

Reactor volume (m ) Mass of catalyst (ton)

REA-1

REA-2

REA-3

REA-4

15.60 22.00

8.32 5.87

7.71 5.44

7.47 5.27

multi- tubular adiabatic reactors. 2.1.5. SNG upgrading unit The raw-SNG, containing methane, CO2, water vapor and H2, is then upgraded to be injected into the grid where it can be stored, transported, or used as fuel (e.g. for CNG cars). Even though the flash unit removed most of the water (see Fig. 4), the raw-SNG must be further dehydrated to comply with the water dew point requirements of the pipelines, to control corrosion and prevent formation of solid hydrocarbon/water hydrates. Currently, glycol dehydrators are widely used for natural gas dehydration. Despite their advantages, the system operation is complex and requires solvent storage, replacement and disposal. Moreover, the vent streams from the dehydrators are a source of emissions of volatile organic compounds. According to Lin et al., membrane technology appears to be an attractive alternative for natural gas dehydration [62]. To this extent, a commercial Pebax®-based membrane (poly ethylene oxide- based block copolymer) was considered to dry the raw-SNG [63,64]. The performance is based on literature values, using an estimated normalized permeability constant of 1000 GPU (1 GPU being 3.35 10−10 mol/m2 s Pa) [62,65], to determine the size required using Eq.(8). In the Aspen Plus™ model, only the retentate/permeate pressures are specified for given feed conditions, as well as the split fractions of one of the stream. The feed gas mixture is compressed in a multi-stage raw-SNG compression step to reach a pressure of 35 bar. A small one-stage system removes 90% of the water in the feed gas, producing a lowpressure permeate gas representing 5–6% of the initial gas flow [66]. The permeate pressure containing the water is set to 1.1 bar [62,65].

Amemb =

permeate flow rate permeance × driving force

37.8–56.5 < 6 vol% < 5 vol% Technically free Traces (ppm)

1 bar. Almost all H2 along with a certain amount of methane transport across the membrane, was recirculated back to the reactor. The large amount of methane not crossing the membrane forms the retentate at the identical pressure as the feed stream. The product stream SNG was then supplied to the natural gas grid usually at elevated pressures from 4 to 70 bar, by using a multistage compressor [68]. The number of stages was determined by a pressure ratio of 4, with equal pressure ratio among all stages. The isentropic efficiency of each stage of 75%, and the inter-stage cooling temperature was set at 40 °C. In the frame of this study, it was assumed that the SNG was delivered at 50 bar. Gas specification for gas grid injection in Germany is presented in Table 5. The Wobbe index represents the interchangeability of fuel gases with respect to natural gas, which is usually used to compare the combustion energy output of different composition fuels.

Table 4 Main specifications for the four multi- tubular adiabatic reactors at a GHSV of 4000 h−1. Parameter

Specification 3

2.1.6. Heat recovery and utilities Systematic process-to-process heat recovery through data evaluation, pinch analysis, and optimized heat exchanger network were performed on the system to reach high energy efficiency and minimize utility costs. Appropriate heat exchangers areas were designed using Aspen Plus™. According to the present design, HX-1 was fed with pressurized boiled feed water at 110 °C and 125 bar. HX-2 was placed to extract heat from the outlet of the first reactor (REA-1). The steam temperature and pressure at the outlet of HX-2 was 505 °C and 125 bar, respectively. A steam turbine unit was used to generate process power to cover a portion of the electrical demand. This was expanded to 10 bar with an isentropic efficiency of 0.85, generating 1055 kWe . After the exit of the turbine, the steam was fed to the condenser, where it released heat and was condensed to 110 °C. The condensate returned to the system with a flow rate equal to 7.5 t/h. As previously mentioned, five additional heat exchangers (from HX3 to HX-7) were implemented to cool down the outlet of each reactor to the working temperature of 350 °C, to obtain high CO2 conversion. They were fed with boiled feed water at 110 °C and 8 bar, to generate a medium-pressured steam at 170 °C and 8 bar. This was done so that it could be brought to the reboiler heat duty of the CO2 capture unit, which requires heat at 140 °C and 5 bar.

(8)

2.2. Techno-economic analysis

where the permeate molar flux (mol/s) through the membrane is driven by the partial pressure difference (Pa) of the component of interest H2O, namely the driving force, and the permeance is expressed as the ratio of the permeability with the thickness of membrane, leading to the normalized permeability constant (mol/m2 s Pa) taken to be a measure of the membrane’s ability to permeate gas. In addition, excess hydrogen also has to be separated in a gas upgrading downstream of the dehydration membrane, and then fed back to the methanation reaction. A commercial membrane technology (polysolfone-based membrane) for hydrogen separation was chosen to achieve a 90% hydrogen separation [67]. An estimated normalized permeability constant of 100 GPU [66] was considered to approximate the performance and determine the size required using Eq. (8). The stream is therefore expanded to reach a working pressure of 20 bar. The permeate pressure containing the component of interest H2 was set to

2.2.1. Technical analysis Technical indicators, including the mass balance of individual inputs and outputs, and the utilities demand (mainly heat and electricity duties), are direct results from the process modelling using the software Aspen Hysys™ and Aspen Plus™. In addition, the methanation efficiency is calculated according to Eq. (9) using the general approach suggested by Salomone et al. [9]: CH4

=

CH4

× XCO2 × MCH4 × LHVCH4 H2 × MH2 × LHVH2

(9)

where XCO2 is the conversion of CO2 within the methanation unit, MCH4 and MH2 are the CH4 and H2 molar weights (kg/kmol), LHVCH4 and LHVH2 are the CH4 and H2 lower heating values (MJ/kg), CH4 and H2 are the CH4 and H2 stoichiometric parameters, respectively. The 7

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conversion of CO2 within the methanation unit is assumed to be superior to 98.9% to achieve the SNG quality [9]. To evaluate the process efficiency of the PtG plant, the PEM electrolyzer, methanation unit, compressors, and heater are considered, according to Eq. (10) [9]: PtG

=

nCH4 × MCH4 × LHVCH4 EPEM + Ecomp + Eheat

The total product cost (TPC ) is calculated by Eq.(14).

TPC =

3.1.1. CO2 capture unit results Table 8 summarizes the CO2 capture unit simulation results in terms of gas phase compositions. An absorption ratio of 90 mol.% leads to the production of 211 kmol/h of CO2 (9.16 t/h) with 98 mol.% purity. In addition to nitrogen (123 ppm) and oxygen (29 ppm), the produced CO2 flow still contains some water (around 2 mol.%) and some other components initially present in the flue gas, namely SO2 (< 50 ppm), NO (< 5 ppm), NO2 (1 ppm) and CO (< 0.5 ppm). As a result of the water-wash sections, the amines (MDEA and PZ) are only present as traces in the washed-gas and the produced CO2 flow. Concerning the energy consumptions, the main energy consuming unit is the reboiler of the regeneration column (5.75 MWth ), while the other thermal demands are for cooling (mainly condenser cooling energy, −1.45 MWth , and lean solution cooling, −2.80 MWth ). The cooling energy of the absorber intercooling is less significant (−0.72 MWth ) and the water-wash cooling energy is very low (−0.01 MWth ). Regarding the electrical demands, while the total pumping energy is low (0.04 MWe ), the electrical consumption of the compressor used in the RVC configuration is not negligible (0.91 MWe ). These results were obtained for the optimum operating parameters given in Table 9, where the specific energy consumption per ton of captured CO2 is also provided. The specific solvent regeneration energy (2.28 GJ/tCO2) is drastically reduced in comparison with conventional CO2 capture processes using MEA 30 wt% as a solvent (generally estimated around 3.5 GJ/ tCO2 [26] and being precisely calculated at 3.36 GJ/tCO2 using the present simulation model), which highlights the interest of using an advanced CO2 capture process such as the one considered in the present study. All the simulation results in relation with the optimization of the CO2 capture unit are provided in the Supporting Information (see SI.1), where comparisons are also performed in terms of equivalent works and utilities costs both for conventional process configurations (without ICA) and for three different solvents (MEA 30 wt%, PZ 40 wt% and MDEA 10 wt% + PZ 30 wt%).

(11)

2.2.2. Economic analysis Both operating and investment costs were evaluated. Most of the capital investment for the equipment was generated directly via AspenTech Economic Analyzer™. In addition, the cost of several individual pieces of equipment was approximated when the cost of a similar item of a different size or capacity was known, using expression (12). The costs were estimated with a nominal accuracy of ± 30%.

S S0

(12)

where the incremental cost C decreases with larger capacities S , C0 and S0 being the reference characteristics values. The exponent takes the economy of scale effect into account, which is often approximated to 0.6. The results were then updated from the original cost index to the year 2018 using the annual chemical engineering plant cost index CEPCI (CEPCI2018 = 603.1) to account for the price development effect (Chemical Engineering magazine for latest values). The total capital investment (TCI ) was determined using the factorial methodology, according to the recommended ratio factors for fluid processing plant by Peter and Timmerhaus [71], and the capital investment (IE ) of the equipment (see Eq. (13)).

(

n

TCI = IE × 1 +

i=1

RFi

)

(13)

where RF is the ratio factor for direct, indirect and working capital, i being the items summarized in Table 6.

3.1.2. CO2 methanation and SNG upgrading units’ results The CO2 methanation unit simulation results in terms of outlet gas compositions are summarized in Table 10. Based on the simulations and the present process design, material and energy streams were estimated and summarized in Table 11. In order to meet the pipeline quality standards, the CH4 mole fraction in the raw SNG of about 70% has to be increased to more than 90% in order to meet the natural gas grid requirements. Table 12 presents the technical features of the utilized membranes. A membrane’s size of 62 m2 was therefore required to achieve a 90% water separation, using Eq. (8). To achieve a 90% H2 separation, the size of the second membrane was estimated to 244 m2. The electrical consumption of the compressor used to reach the working pressure of the dehydration membrane is of importance (0.36 MWe ). In addition, the SNG was brought to 50 bar to meet the pressure pipelines, leading to a consumption of 0.25 MWe . It is worth mentioning that for some local or regional grid injections, this elevated pressure is not necessary, as these pipelines mostly operate in the pressure range of 4 to 40 bar.

Table 6 Estimation of the total capital investment (TCI ) [71,72]. Type

Item (i )

Direct cost

1 2 3 4 5 6 7 8

Purchased equipment (delivered) Purchased equipment installation Piping Instrumentation and Controls Electrical systems Buildings (including services) Yard improvements Service facilities

1.00 0.47 0.36 0.68 0.11 0.18 0.10 0.70

Indirect cost

9 10 11 12 13 14

Engineering and supervision Construction expenses Legal expenses Contractor’s fee Contingency Working Capital

0.33 0.41 0.04 0.22 0.44 0.89

(14)

3.1. Process performances

The following parameters are also required for the produced SNG stream: (i) composition, essential to achieve the gas requirements for injection into the networks (i.e. percentage of hydrogen); (ii) the Low Heating Value (LHV) and/or High Heating Value (HHV); and (iii) the Wobbe index.

C = C0

Cj

3. Results and discussions

(10)

Eheat EPEM + Ecomp

=1

j=1

with Cj the cost of item j based on Table 7.

where EPEM is the electrical energy used by the electrolyzer (MJ), Ecomp the energy spent for the compression (MJ), and Eheat represents the total required heat (MJ). Finally, the overall energy utilization factor is calculated with Eq. (11). PtG

n

Ratio Factor (RF )

8

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Table 7 Economic assumptions for operating cost estimation (TPC ). Item ( j )

Assumptions

1

Raw material

2 3

Utilities Operation and maintenance

Solvent: 1 €/kg(1); Catalyst: 90 €/kg(2); Process water: 1 €/ton(3); Demineralized water: 3 €/ton(4); Renewable H2: onsite production(5); SNG selling price: 21 €/ton(6); O2 selling price: 87.4 €/ton(7); CO2 credit tax: 25.2 €/ton(8) Cooling water: 0.025 €/ton(9); Electricity: 40 €/MWh(10); Steam: self-produced(11) 60 labors/shift, 3 shift/day, 54,000 € per labor per year(12) + 20% of operating labor

4 5

Depreciation General expenses

Operating labor and supervisory labor Maintenance and repairs Operating supplies Laboratory charges

6% of capital investment 15% of Maintenance and repairs 15% of operating labor Annual depreciation cost, with an assumption of a 20-year recovery period and 0 € salvage value, linear 15% of total product cost

(1) Amine solvent considered: PZ 30% + MDEA 10% (+60% water): 2.16 €/kg PZ × 0.3 + 1.7 €/kg MDEA × 0.1 + 0.001 €/kg H2O × 0.6 = 0.818 €/kg ≈ 1 €/kg solvent cost estimated from de Medeiros et al. [73]. (2) Catalyst cost considered from own estimation. (3) Process water cost considered from Michailos et al. [74]. (4) Demineralized water cost considered from Li et al. [75]. (5) Renewable H2 considered produced onsite. Only the cost of demineralized water is taken into account in the calculation of TPC . (6) SNG selling price estimated from natural gas price (≈ 4.56 €/1000 ft3 in April 2019). (7) Oxygen selling price from Michailos et al. [74]. (8) CO2 credit tax (€/ton) in April 2019, according to European Emission Allowances (EUA). (9) Cooling water cost considered from Michailos et al. [74]. (10) Electricity considered from renewable energy (wind power). Cost considered from Glenk and Reichelstein [76] supported by www.eex.com. (11) Steam considered self-produced. Only the cost of process water is taken into account in the calculation of TPC . (12) Operating labor and supervisory labor costs considered from Zhang et al. [72].

Table 8 Composition of the gas to be treated, washed-gas and produced CO2 (mol frac. basis). Component

Gas to be treated

Washed-gas

Produced CO2

N2 CO2 O2 H2O CO SO2 NO NO2 MDEA PZ

64.7% 20.4% 8.6% 6.2% 1330 ppm 111 ppm 474 ppm 2 ppm 0 0

69.9% 2.2% 9.3% 18.5% 100 ppm 112 ppm 512 ppm 2 ppm Traces Traces

123 ppm 98.0% 29 ppm 1.98% 0.4 ppm 43 ppm 3 ppm 1 ppm Traces Traces

Total mol flow (kmol/h) Total mass flow (t/h)

1124 34.85

1040 27.90

211 9.16

Table 10 Composition of the outlet streams (mol frac. Basis).

Specification

Operating conditions

(L/G) vol,opt (m3/m3) L opt (m3/h) Intercooling temp. (°C) Intercooling stage Flash p (bar) CO 2, rich (mol/mol) CO 2, lean (mol/mol)

Energy consumption

Eregen (GJ/tCO2)

Inlet

REA-1

REA-2

REA-3

REA-4

Raw-SNG

SNG

CO2 H2 H2O CH4 CO

19.9% 79.9% 0.2% 0% 0%

7.5% 31.8% 40.5% 20.2% Traces

4.3% 19.0% 51.1% 25.6% –

2.4% 11.5% 57.4% 28.7% –

1.5% 8.1% 60.3% 30.2% –

3.7% 21.3% 0.3% 74.8% –

3.7% 3.4% Traces 92.9% –

Total mol flow (kmol/ h) Total mass flow (t/ h)

1038.12

745.23

692.67

664.94

653.05

268.54

214.22

10.84

10.79

10.79

10.79

10.79

3.78

3.60

Table 11 Material and energy consumptions of the PtG plant.

Table 9 Optimum operating parameters and specific energy consumptions of the CO2 capture unit. Indicator

Component

Indicator

Specification

Value

Value

CO2 methanation inputs

5.19 10−3

Energy consumptions

40 8 4 0.98 0.33

H2 (t/h) CO2 (t/h) CO2 compressor (MWe ) Recycle compressors (MWe ) Membrane compressor (MWe ) SNG compressor (MWe ) Pump (MWe ) Turbine (MWe )

1.65 9.09 0.14 0.19 0.36 0.25 0.07 −1.06

PtG outputs

SNG (t/h) SNG Wobbe Index (MJ/m3) Total Output HHV (MW)

3.60 46.72 50.4

Methanation efficiency (

83.4

128

2.28 Efficiency

3.1.3. Process integration results The methanation efficiency was estimated at 83.4%, while the process efficiency of the PtG plant ( PtG ) equaled 65.5% (72.6% on HHV basis), based on Eq.(9) and (10), respectively, with a comprehensive usage of excess heat and technical assumptions. The overall energy utilization factor ( PtG ) reached 75.7%. A Wobbe Index of 46.72 MJ/m3 was calculated, which is in the range of Wobbe Indices in Germany (see

CH4 )

PtG plant HHV ( PtG ) Overall energy utilization factor (

PtG )

72.6 75.7

Table 5). These results are summarized in Table 11. Stream data (temperature, composition, heat capacities, etc.) computed by Aspen Plus™ are available in the Supporting Information (see SI.4). A minimum pinch point temperature of 30 °C was defined for the 9

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the total product cost, it was estimated at 61.60 M€ per year. Fig. 6 displays the operational expenses distribution, based on the previously mentioned assumptions (see Table 7). The expected indirect revenues were 407.9 € per ton SNG, considering SNG, oxygen and CO2 trading. At present, the estimated cost of SNG is therefore 1587 €/ton SNG, about 115 €/MWh, which is in the cost range evaluated by Guilera et al. [2]. It is about 3.5 times higher than fossil natural gas (33 €/MWh in Germany [2]). The annual operating costs amount to around 7% of the investment, which is in accordance with the values reported for catalytic processes [78]. It is worth noting that the amortized CAPEX of the electrolyzers is taken into account in “Hydrogen production”. A sensitivity analysis was carried out regarding the parameters that most influence the SNG cost, derived from Table 15. One input variable at a time was changed to its low or high value, while the other variables were maintained at their base values. Fig. 7 illustrates the sensitivity analysis. The electricity price therefore has a huge influence on the cost of SNG, especially due to the fact that H2 production derived directly from it, represents about 70% of the cost of SNG. The additional expenses and CAPEX, the electrolyzer efficiency, and the O2 selling price have an influence higher than 10% on the final SNG cost. In particular, the income obtained from the sale of oxygen in the PtG plant represents an interesting economic incentive, especially in the current context of low prices for carbon emissions trading (CO2 credit tax) that have an almost negligible effect.

Table 12 Technical features of the membranes system. Parameter 2

Membrane area (m ) Feed pressure (bar) Permeate pressure (bar) Temperature (°C) Permeance (GPU) Compressor (MWe )

Dehydration membrane

H2 separation membrane

62 35 1.1 40 1000 0.36

244 20 1 90 100 –

Temperature (°C)

600 500 400 300 200 100 0

0

2

4

6

8

10

12

14

Heat load (MW) Hot composite curve

Cold composite curve

Fig. 5. Hot and cold composite curves.

4. Conclusion and outlook

heat exchanger with gaseous streams [77]. The total recovered heat from the plant amounts to about 13.25 MWth , which is available between 25 and 615 °C (see Fig. 5). A fraction of the heat recovered at 170 °C can be brought to the reboiler of the CO2 capture unit, accounting for 5.75 MWth , so that the heat integration compensates the hot utility requirement of the CO2 capture unit. In addition, the electricity produced in the steam turbine was used to cover internal demand (see Table 11). Thus, no external electrical requirement was required for both the CO2 methanation and SNG upgrading units. The total external electrical requirement of the whole suggested process designed is therefore linked to the compressor used in the RVC configuration, equal to 0.90 MWe .

The present work has assessed the production of SNG from industrial CO2 in the context of PtG technology. The PtG plant was implemented in AspenTech’s software, both in Aspen HYSYS™ and Aspen Plus™, to extract mass and energy balances in order to investigate its economic viability. An advanced CO2 capture process was proposed to treat 10% of a BAT cement plant’s flue gas. Together with renewable H2, they were converted to SNG through methanation reaction, producing 0.40 ton SNG per ton of captured CO2. In the suggested process design, a dry methane mole fraction of 92.9 mol.% in the final SNG was achieved, together with a CO2 and H2 content of 3.7 mol.% and 3.4 mol. %, respectively, which are the most critical aspects of the SNG quality and compatibility. Systematic process-to-process analysis was investigated to reach high energy efficiency and minimize utility costs. Thermal integration led to an overall system HHV-based efficiency equal to 72.6%. As seen in this work, SNG is currently not competitive with natural gas. The estimated cost of SNG was about 115 €/MWh, about 3.5 times higher than fossil natural gas. Costs are mainly driven by the high initial investment costs, as well as the production of H2, which depends directly on the price of renewable electricity. Thus, a reduction of the electrolyzer investment costs combined with a higher electrolyzer efficiency and higher revenues, could result in a profitable business case for large-scale deployment of this PtG technology. In future scenarios, the cost of SNG can be therefore reduced. Specific costs for PEM technologies are for instance expected to fall from about 1200 €/kW in 2017 to 500 €/kW in 2050 [40]. Furthermore, when applying a time-ofuse retail electricity price of approximately 10 €/MWh, the costs of SNG decrease to about 40 €/MWh, which is really close to the current price of fossil natural gas. Several studies also point out that with ongoing trends in 10–20 years, PtG might be economically profitable [33]. Moreover, in regions where a natural gas infrastructure exists, the existing network for grid gas injection is a major advantage of the PtG technology. By utilizing excess energy to produce hydrogen via water electrolysis, energy can be stored and distributed in the existing system for use when and where it is needed. Despite the fact that the use of (excess) renewable energy introduces a seasonal and daily component into H2 production, Gahleitner highlighted that the use of such fluctuating power sources is satisfying, and is currently in use in a lot of PtG

3.2. Economic performances Based on the simulation results, the capital investment and operating costs of the integrated process plant were evaluated. Table 13 lists the estimated capital investment for the equipment. Table 14 presents the main costs for all the units. The TCI of the integrated process was therefore evaluated to 166.85 M€ (excluding the capital investment of H2 production), where the CO2 conversion unit itself represents 58%, mainly due to the complex network of heat exchangers that had been implemented to integrate the heat. The final TCI was estimated at 804.92 M€, where about 79% was dedicated to the renewable production of H2. Regarding Table 13 Cost estimation of main equipment (M€). Main equipment

CO2 Capture

CO2 methanation

Gas Upgrading

Columns (absorber/stripper) Reactors Compressors Turbine Heat Exchangers (heaters/ coolers) Flash tanks Membranes Total Purchased equipment cost (delivered)

1.24 – 3.36 – 2.40

– 3.50 2.28 0.45 9.35

– – 3.91 – 0.15

0.18 – 7.18

0.09 – 16.22

– 0.68 4.74

10

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Table 14 Estimation of the total capital investment (M€). Main cost

H2 production (1)

Total Direct Cost Total Indirect Cost Fixed Capital Investment (FCI ) Total Capital Investment (TCI )

387.36 154.94 542.30 638.07

CO2 Capture

CO2 methanation

Gas Upgrading

25.84 10.34 36.17 42.56

58.40 23.36 81.76 96.20

17.05 6.82 23.88 28.09

(1) Calculation based on specific electrical power of approximatively 55.6 kWh per kg H2 and an investment cost of 1200 € per kW (installed) for the production of 40 ton per day of renewable H2 [49,76], and Table 6 to estimate the Total Direct Cost.

Fig. 6. Operational expenses in € per ton SNG.

targets, and high renewable energy availability and penetration in the market. Additionally, several actions could be taken to foster the deployment of PtG technologies. A key action is the development and deployment of cheap low carbon electricity, which is necessary to reduce the costs for PtG. The transition to low-carbon depends on the cost for conventional fossil choices, such as the gas price. Thus, additional taxes on fossil resources could promote a shift to PtG technologies. Direct subsides is also a measure that can be used to improve the business case for private investors. Creating a green hydrogen market would also be necessary, giving a premium in the selling price of hydrogen. Finally, to overcome the seasonal component of PtG, nuclear, geothermal, biomass, and hydro energy are necessary, also increasing the flexibility of the whole system. By applying these actions, combined with a policy penalizing new investment in industries using fossil resources, the suggested process will likely play a certain role for the transition to a renewable network, integrating large fractions of renewables that require balancing power and seasonal energy storage, and introducing considerable flexibility into the energy system. Furthermore, additional energy and environmental policy challenges can be solved by the enlargement of the percentage of alternative fuels in the mobility or heating sector. In a longer-term perspective, with markets characterized by cheap electricity and high fuel prices, PtG shall gradually eliminate fossil fuels, transitioning to a low carbon economy. PtG may also have higher acceptance in society, denoting advantages over the existing alternative storage technologies as well as over network expansion, representing a non-negligible component in the future prospects of the technology.

Table 15 Input variable for the sensitivity analysis. Input variables (1)

Electricity price (€/MWh) Electrolyser efficiency(2) (%) Additional expenses & CAPEX(3) CO2 credit tax(4) (€/ton) Oxygen price(5) (€/ton)

Low value

Base value

High value

0 65 −30% 0 43.7

40 75 Nominal 25.2 87.4

120 85 +30% 50.4 131.1

(1) Electricity ranging from 0 €/MWh (free renewable electricity) to 120 €/MWh (own assumption). (2) Electrolyser efficiency commonly considered in the literature [74,76]. (3) Additional expenses and CAPEX accuracy with a ± 30% fluctuation. (4) Electricity considered CO2 credit tax in April 2019, according to European Emission Allowances (EUA). (5) Oxygen selling price from Michailos et al. [74], with a ± 50% fluctuation.

Electricity price Additonal expenses & CAPEX Electrolyser efficiency O₂ selling price CO₂ credit tax 0

50

Low value

100 150 200 SNG cost (€/MWh)

250

300

High value

Fig. 7. Sensitivity analysis on the cost of SNG in €/MWh (base cost 115 €/MWh SNG).

CRediT authorship contribution statement

pilot plants [79]. In the context of CCUS, the main drivers for methanation deployment are limited and/or unwanted CO2 storage, high CO2 reduction

Remi Chauvy: Conceptualization, Methodology, Investigation, Validation, Writing - original draft, Writing - review & editing, Visualization. Lionel Dubois: Methodology, Investigation, Validation, Writing - original draft, Writing - review & editing. Paul Lybaert: 11

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Resources, Validation. Diane Thomas: Writing - review & editing, Supervision. Guy De Weireld: Validation, Writing - review & editing, Supervision.

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