Power-to-Green Methanol via CO₂ Hydrogenation—A Concept Study including Oxyfuel Fluidized Bed Combustion of Biomass

Abstract

A revolution of the global energy industry is without an alternative to solving the climate crisis. However, renewable energy sources typically show significant seasonal and daily fluctuations. This paper provides a system concept model of a decentralized power-to-green methanol plant consisting of a biomass heating plant with a thermal input of 20 MW_th. (oxyfuel or air mode), a CO₂ processing unit (DeOxo reactor or MEA absorption), an alkaline electrolyzer, a methanol synthesis unit, an air separation unit and a wind park. Applying oxyfuel combustion has the potential to directly utilize O₂ generated by the electrolyzer, which was analyzed by varying critical model parameters. A major objective was to determine whether applying oxyfuel combustion has a positive impact on the plant’s power-to-liquid (PtL) efficiency rate. For cases utilizing more than 70% of CO₂ generated by the combustion, the oxyfuel’s O₂ demand is fully covered by the electrolyzer, making oxyfuel a viable option for large scale applications. Conventional air combustion is recommended for small wind parks and scenarios using surplus electricity. Maximum PtL efficiencies of η_PtL,Oxy = 51.91% and η_PtL,Air = 54.21% can be realized. Additionally, a case study for one year of operation has been conducted yielding an annual output of about 17,000 t/a methanol and 100 GWh_th./a thermal energy for an input of 50,500 t/a woodchips and a wind park size of 36 MWp.

1. Introduction

The climate crisis poses a major threat to human civilization on a long-term basis. The global monthly average CO₂ concentration in the atmosphere reached a level of 415 ppm in December 2020, with average growth rates of 2–3 ppm per year in the previous five years [1,2]. Combustion processes for power generation account for about 42% of global anthropogenic CO₂ emissions, indicating the energy industry’s key role for mitigating global warming [3]. In 2019, the European Commission launched its Green Deal program, including the goal to become the first climate-neutral continent by 2050, as well as decoupling economic growth and resource usage [4].

To reach these ambitious targets, intersectional concepts, based on renewable energy sources, for the provision of power and heat need to be empowered and improved. Carbon capture and storage (CCS) and carbon capture and utilization (CCU) pose promising technologies in fighting the climate crisis. However, CCS goes hand in hand with several difficulties (e.g., safety issues and substantial efforts for the transportation of captured CO₂). Additionally, CCS is illegal in numerous countries and therefore, a limited option for the mitigation of CO₂ emissions. A variety of promising CCU technologies (e.g., power-to-gas (PtG), power-to-liquid (PtL), power-to-fuels (PtF) and power-to-chemicals (PtC)) are on the verge of becoming economically feasible within this decade. Significant advantages are the on-site usage of CO₂ sources, as well as the development of new business cases to produce renewable fuels or chemicals. Crucial parameters defining the sustainability of CCU technologies are the CO₂’s origin as well as the product’s (H₂, CH₄, methanol, Fischer-Tropsch products) scope of application. Possible CO₂ capture systems (i.e., direct air capture (DAC), post combustion, pre combustion, oxyfuel combustion or industrial processes) are listed and explained in detail in the IPCC special report on carbon dioxide capture and storage [5]. In general, sources with a high concentration of CO₂ are favored due to a smaller technological effort of CO₂ capture.

In May 2018, 128 PtG projects were in progress or have been finished in Europe. A list of conducted projects as well as an overview of possible PtG process routes can be found in [6]. Three PtG demo sites for CO₂ valorization to methane using different reactor technologies (i.e., a catalytic honeycomb reactor, a biological stirred bubble column reactor and a catalytic milli-structured reactor) are analyzed in [7]. Within the pan-European project MefCO₂, a PtL pilot plant with an annual output of 500 t/a methanol was realized. Captured CO₂ was valorized using a polymer electrolyte membrane electrolyzer powered with solar-generated electricity in Niederaußem, Germany [8]. Another successful example is the George Olah Plant, located in Iceland, a PtL plant at industrial scale, designed and constructed by the Icelandic company Carbon Recycling International (CRI). Over 4000 t/a methanol can be produced valorizing geothermal CO₂ with H₂ generated by an alkaline electrolyzer powered by Iceland’s 100% renewable grid electricity [9]. After winning the Nobel Prize in 1994, George Olah’s scientific focus shifted towards producing green fuels (e.g., methanol using captured CO₂ as feedstock), making him a pioneer in PtL processes before the term came into existence [10]. The German start-up company, INERATEC, announced the construction of a PtF plant at industrial scale at Frankfurt, Germany, until 2022. A total of 3500 t/a Fischer-Tropsch products will be produced using a maximum of 10,000 t/a biogenic CO₂ [11]. Research analyzing CO₂ hydrogenation to methanol at lab-scale can be found in [12] (applying an In₂O₃/ZrO₂ catalyst) and [13] (conventional Cu/ZnO/Al₂O₃ catalyst). Within the past years PtL technologies transitioned from a niche application at laboratory scale to a viable option for the process and energy industry regarding CCU technologies. Current projects aim at proofing the process’ feasibility at pilot scale obtaining reasonable efficiency rates. Major challenges within the following years will be the provision of cost-effective renewable H₂ by electrolyzers surpassing a power input of several MW. Power-to-liquid processes have the potential to provide renewable eFuels as the intermediate power source until individual mobility is fully electrified. In the long run, eFuels could be applied as the power source for aviation and goods transport, requiring a high energy density.

Oxyfuel combustion was first proposed in 1982 to provide a CO₂-rich gas stream for enhanced oil recovery applications, and experienced a technological renaissance in the early 2000s when being rediscovered as a suitable option for CCU processes. Oxyfuel combustion tackles the issue of low CO₂ concentrations in flue gases obtained by conventional air combustion processes [14]. In oxyfuel combustion technologies, fuels are burned in a mixture of O₂ with a technical purity of at least 95 vol% and recirculated flue gas. The recirculation of flue gas is mandatory, as the missing volume stream of the air’s N₂ needs to be replaced. Furthermore, the combustor’s flame temperature can be controlled by adjusting the amount of recirculated flue gas. The main differences when comparing oxyfuel combustion with conventional air combustion are [14]:

• O₂ concentrations of about 30 vol% in the oxidant are required to obtain a comparable flame temperature;
• Volume flow streams of emitted flue gas are reduced by about 80% since no N₂ passes the combustor as inert gas;
• Different combustion regimes are observed due to deviating physical properties of CO₂ and N₂. For example, the density of flue gas generated by an oxyfuel combustion process is larger due to a higher molar mass of CO₂, concluding in a larger specific heat capacity as well.

The “Callide” oxyfuel project, successfully completed in 2015, can be considered a showcase regarding oxyfuel combustion at an industrial scale. A power plant located in the state of Queensland, Australia, with a thermal input of 24–29 MW_th. (coal) was combined with a CCS project located in Victoria, Australia. Detailed scientifical results can be found in the project’s final report [15]. Another oxyfuel project launched by Vattenfall in 2006 at Schwarze Pumpe, Germany, aimed at retrofitting an existing coal power plant with a thermal input of 30 MW_th. for CCS applications. However, Vattenfall cancelled the project in 2014 due to economic difficulties without publishing a final report. A detailed study on oxyfuel combustion of biomass in a 20 kW_th fluidized bed combustor can be found in [16], analyzing effects of combustion environments (oxyfuel combustion vs. conventional air combustion) on fuel combustion, temperature profiles and CO and NO_x emissions.

As stated previously, power-to-X technologies will play a significant role in fighting the climate crisis due to their potential to chemically store excess power generated by renewable energy carriers as H₂, CH₄, methanol or Fischer-Tropsch products. Renewable power sources (i.e., wind or solar) undergo daily and seasonal fluctuations leading to difficulties regarding the storage of surplus electricity. Generated energy carriers could also be integrated in grid-stabilizing systems, tackling another major disadvantage of renewable power sources.

Furthermore, hydrogenating CO₂ comes with several process advantages compared to conventional methanol production by syngas [9,17]:

• Fewer impurities can be found in the crude methanol;
• The chemical reaction of CO₂ and H₂ is less complex and less exothermal compared to the synthesis based on CO and H₂;
• Boiling water reactors (BWRs) including a recirculation of unconverted gases can easily be applied instead of adiabatic reactors in series since less reaction heat needs to be transferred out of the catalyst bed, resulting in a larger economic feasibility of the process;
• Milder process conditions are required;
• The methanol selectivity of CO₂ hydrogenation is larger compared to conventional methanol synthesis processes based on syngas as the reactor input.

The design and simulation, using Aspen Plus, of a plant producing methanol via CO₂ hydrogenation can be found in [18]. CO₂ generated by a thermal power plant using coal as fuel was hydrogenated with H₂ produced by water electrolysis. The authors declared that O₂, generated by the electrolyzer, had to be sold as a by-product and highlighted the possible potential of using it as feedstock for oxyfuel combustion processes without calculating detailed scenarios. A feasibility analysis of a plant producing renewable methanol by chemically storing wind power can be found in [19]. The production costs of H₂, being the most significant model parameter, as well as the selling price of methanol were varied. Consequently, selling or using O₂ generated by water splitting could be a major factor in improving the facility’s economic feasibility.

This article aims at finding and conceptualizing different process routes for a decentralized power-to-green methanol plant, with a thermal input of 20 MW_th., using woodchips as fuel (Figure 1). As a result, the proposed concept has the potential to achieve a negative CO₂ balance. A foundation for future research projects should be provided with the goal to increase the concept’s level of details. Furthermore, proposed process routes will be evaluated by comparing obtained power-to-liquid efficiency rates to values stated in reviewed literature. In comparison with previous studies, this paper includes the application of oxyfuel combustion in combination with methanol production (CO₂ hydrogenation) by directly applying O₂ (by-product of water splitting) to the oxyfuel combustor. Furthermore, a direct comparison of oxyfuel and air process routes in combination with CCU, regarding the efficiency rates of power-to-green methanol facilities, should answer the question whether oxyfuel process routes are a viable option for future PtL plants.

A specific hypothesis of this work is whether the replacement of conventional air by oxyfuel combustion enhances the performance of power-to-methanol plants using post combustion CO₂ capture. The preferable combustion technology as well as the synergy between oxyfuel combustion and water electrolysis (generating O₂ as byproduct) was determined for different assumed scenarios.

2. Materials and Methods

The main aspect of this work was to model several subprocesses (oxyfuel combustion in combination with a DeOxo reactor, air combustion including CO₂ capture by absorption, methanol synthesis, alkaline electrolyzer). Combining them resulted in models of different process routes. Microsoft Visio (2016) was applied for process visualization (i.e., creating process flow charts). A list of auxiliary equipment for all subprocesses can be found in the Appendix A (

Table A9. Machine list.

Name	Machine
H1	Twin screw conveyor
V1	Blower
V2,V3	Fans
W1	Preheating heat exchanger
W2–7	Heat exchangers
C1–4	Multi-staged compressors
P1–8	Rotary pumps
D1,D2	Relief valves
F1,F2,F3,F4	Distribution valves

).

2.1. Oxyfuel Combustion in a Bubbling Fluidized Bed Combustor (BFBC)

As described previously, oxyfuel combustion is defined as combustion reaction in a mixture of O₂ with a technical purity of at least 95 vol% and recirculated flue gas. A standard combustion calculation was conducted.

Fuel-N and fuel-O were assumed to be released as N₂ and O₂, respectively. Gases were assumed to behave like ideal gases. Complete combustion was assumed for subsequent combustion calculations. A similar approach can be found in [20] for comparison.

Subsequent mass balances are based on assumed properties of woodchips stated in

Table 1. Assumed ultimate analysis of woodchips.

Ultimate Analysis—Woodchips (Dry and Ash Free) [wt%]
C	H	N	S	O	SUM
50	6	0.4	0.05	43.55	100

(ultimate analysis) and

Table 2. Assumed proximate analysis of woodchips.

Proximate Analysis—Woodchips [wt%]
Water Content W	Ash Content A	Volatiles	Fixed Carbon
30	2	80	20

(proximate analysis). These parameters serve as a benchmark but can easily be adjusted to different scenarios or alternating fuels (e.g., bark or municipal waste). The fuel’s lower heating value (LHV) in MJ/kg was calculated with the equation after Boie. Small letters refer to the different mass participations of the fuel’s elemental components in kg/kg fuel.

LHV = 34.8∙c + 93.9∙h + 6.3∙n + 10.5∙s − 10.8∙o − 2.5∙w [MJ/kg].

(1)

The combustion reaction’s specific O₂ demand per kg of combusted species, stated in the Appendix A (

Table A1. Specific O₂ demand of the combustion reaction per species i.

Species	O2 Demand [Nm3/kg of Species i]
C	1.8659
H	5.5608
O	−0.7003
S	0.6988

), was calculated with previously stated assumptions.

O₂ is either provided by an air separation unit (ASU) or the alkaline electrolyzer. Excess of O₂ in the oxidant α was defined as the ratio of actual volume flow stream of O₂ related to the volume flow stream of O₂ required for a stoichiometric combustion, as described in [20].

$$\alpha =\frac{{\dot{V}}_{Ox.,O_2,Real}}{{\dot{V}}_{O_2,Stoichiometric}}.$$

(2)

Since the recirculation of flue gas is mandatory for controlling the combustion temperature, O₂ provided by the recycled flue gas needs to be considered as well. The excess of O₂ in the fluidization medium, defined as sum of O₂ provided by an ASU or alkaline electrolyzer plus recirculated flue gas, was defined by parameter β, being comparable to the air ratio λ for conventional air combustion calculations.

$$\alpha =\frac{{\dot{V}}_{FM,O_2,Real}}{{\dot{V}}_{O_2,Stoichiometric}}=\frac{{\dot{V}}_{Ox.,O_2,Real}+{\dot{V}}_{FGR,O_2,Real}}{{\dot{V}}_{O_2,Stoichiometric}}.$$

(3)

A definition of the recirculation ratio r, the ratio of volume flow rate of recirculated flue gas related to the volume flow rate of generated flue gas, can be found in [21].

$$r=\frac{{\dot{V}}_{FGR}}{{\dot{V}}_{FG}}.$$

(4)

Condenser modeling was done with the Antoine Equation (5) and water loading X of flue gas for specific temperatures (6). The assumed parameters A, B and C for water are listed in the Appendix A (

Table A6. Parameters for water—Antoine equation.

Parameter	Value [-]
A	17.2799
B	4102.99
C	237.431

). Parameters of water at the triple point (p_Tr. = 611.657 Pa) are stated in [22].

$$\ln \left(\frac{p_S}{p_{Tr.}}\right)=A+\frac{B}{\vartheta +C}.$$

(5)

$$X=\frac{0.622·p_S}{p-p_S}.$$

(6)

Assumed parameters regarding the BFBC’s design, based on benchmarks given in [23] for a thermal input of 20 MW_th., are listed in the Appendix A (

Table A7. Assumed parameters of the BFBC.

Parameter	Value [Unit]
Bed surface area ABed	15 m2
Bed volume VBed	15 m3
Bed height hBed	1 m
Bed diameter dBed	4.37 m
Bed temperature TBed	850 °C

). BFBC is the best available technology for combusting fuels with high contents of volatiles (e.g., biomass) for small-scale decentralized applications [24].

To ensure the fluidization of bed material (e.g., silica sand and woodchips), a fluidization velocity u_f of 1–2.5 m/s is recommended [25]. Mixing inside the reactor is increased significantly with increasing u_f, leading to an improved combustion reaction regime. As a result, segregation of bed material is prevented and the bed’s heat transfer is improved [26].

The amount of utilized CO₂ is a function of H₂ provided by the alkaline electrolyzer, since H₂ is required for CO₂ hydrogenation inside the methanol reactor. In real applications, the generation of H₂ will vary due to fluctuating power levels provided by a wind park. The share of utilized CO₂ was defined by the allocation coefficient γ, describing the ratio of flue gas being transferred to downstream process steps and the total amount of flue gas generated. In case no H₂ is generated by the electrolyzer, the total amount of generated flue gas is released into the atmosphere (γ = 0).

$$\gamma =\frac{{\dot{V}}_{FG to MeOH synthesis}}{{\dot{V}}_{FG, total}}.$$

(7)

2.2. Conventional Air Combustion in a BFBC

Comparing different process routes of a decentralized power-to-green methanol plant, processing CO₂ generated by the combustion of woodchips, was the main goal of this work. Therefore, modeling a scenario using conventional air combustion technology based on assumptions listed in Section 2.1 was mandatory. The following differences were included:

• Ambient air is used as fluidization medium instead of a mixture of O₂ and recirculated flue gas;
• Parameter β was replaced by λ (conventional air ratio);

Physical properties of air can be found in [22]. Components like H₂O, CO₂ and noble gases other than Argon were neglected.

2.3. Flue Gas Cleaning

A simplified flue gas cleaning model was introduced in the plant’s model. Desulfurization is realized by injecting Ca(OH)₂ (i.e., slaked lime) into the gas stream prior to a baghouse filter unit, ensuring fly ash removal. SO₂ and Ca(OH)₂ react to CaSO₃, subsequently collected by the baghouse filter, and water.

$${\mathrm{SO}}_2+\mathrm{Ca}{\left(\mathrm{OH}\right)}_2\to {\mathrm{CaSO}}_3+{\mathrm{H}}_2\mathrm{O}.$$

(8)

The required mass flow rate of Ca(OH)₂ for desulfurization was calculated stoichiometrically based on Equation (8) including an assumed excess of 50% Ca(OH)₂. A desulfurization efficiency of 95% was assumed [27].

Removing NO_x from the flue gas stream is achieved by using a SCR-DeNO_x unit. NH₄OH (i.e., ammonia water), with a concentration of 19–29 wt%, is inserted prior to a ceramic honeycomb catalyst at temperatures in the range of 453–703 K, resulting in the reduction of NO_x to N₂ and water [28].

$$4\mathrm{NO}+4{\mathrm{NH}}_3+{\mathrm{O}}_2\to 4{\mathrm{N}}_2+6{\mathrm{H}}_2\mathrm{O}.$$

(9)

The flue gas’ NO_x concentration was based on experimental data listed in [25]. Wooden pellets with a fuel-N content of 0.3 wt% (dry and ash free) were burned in a BFBC under lab conditions. NO_x emissions are converted to the fuel-N content of the assumed woodchips stated in Section 2.1. The injected amount of NH₄OH was assumed to be 50% larger than stoichiometrically required for the conversion of NO_x according to Equation (9).

2.4. CO₂ Processing

Before entering the methanol synthesis process, generated CO₂ needs to be concentrated and processed. The applied technology depends on the upstream combustion technology (oxyfuel vs. air combustion). Flue gas produced by oxyfuel combustion processes is rich in CO₂ after removing H₂O in a condenser. Therefore, mainly excess O₂ needs to be removed (inert N₂ does not pose a problem in the downstream methanol synthesis process). When using conventional air combustion, CO₂ needs to be captured and concentrated (e.g., realized by a monoethanolamine (MEA) absorption process). Both configurations require a catalyst protection unit, for example, a fixed bed reactor using ZnO or CuO as adsorbent to remove catalyst poisons (e.g., H₂S and SO₂) from the flue gas stream.

2.4.1. DeOxo Reactor

Removing excess O₂ is necessary when using oxyfuel combustion technologies, since β will always exceed a value of 1. The following assumptions were made regarding the modeling of a DeOxo reactor:

• Gases behave like ideal gases;
• The amount of required H₂ was calculated stoichiometrically;
• H₂O and N₂ were considered inert gases and therefore, have no influence on the reactor’s performance;
• Reactor pressure p_DeOxo = 10 bar;
• Heat required for the regeneration of activated alumina was not taken into account;
• Reactor temperature T_DeOxo = 323 K.

Equation (10) describes the reactor’s underlying chemical reaction. The catalyst bed must be cooled since the reaction is strongly exothermic. H₂O is formed as product and needs to be separated by two adsorbent beds loaded with activated alumina. As noted previously, a catalyst guard unit must be installed in front of the DeOxo reactor.

$${\mathrm{H}}_2+\frac{1}{2}{\mathrm{O}}_2\to {\mathrm{H}}_2\mathrm{O} \Delta H_r^0=-285.8\frac{kJ}{mol}.$$

(10)

2.4.2. CO₂ Capture—MEA Absorption

CO₂ capture based on absorption processes using aqueous MEA solutions obtain the highest technology readiness level (TRL) regarding carbon capture technologies [29] and are solely applicable at an industrial scale [30]. Detailed information regarding process parameters and reactor design can be found in [31]. Alternatively, CO₂ capture processes can be realized using adsorption processes in fluidized bed reactors. Experimental data and data obtained from simulations can be found in [32,33,34,35,36,37].

The modeled CO₂ capture system was required for process routes using conventional air combustion with flue gas CO₂ concentrations below 20 vol%, including an absorber, a regenerator, conveyor units cycling the MEA solution and a heat exchanger improving the process’ heat integration. Main model input parameters were the amount of utilized flue gas as well as the flue gas’ composition. The first output stream, a purified CO₂ stream, is transported to a downstream methanol synthesis unit. Off-gas, with a low concentration of CO₂ leaving the regenerator, is added to non-utilized flue gas and emitted to the atmosphere. The following parameters have been assumed:

• CO₂ capture efficiency η_Capture = 0.9;
• Energy demand $Q_{C{O}_2,MEA}^{’}$= 3.8 MJ/kg CO₂ captured;
• Absorber temperature T_Abs. = 308 K;
• Regenerator temperature T_Des. = 393 K;
• Pressure p_Abs. = p_Des. = p_Atm.;
• CO₂ stream purity = 𝑦_CO2 = 99 vol%.

Additionally, subsequent assumptions have been made:

• Gases behave like ideal gases;
• Stripping gas required for the desorption was not considered;
• Examined species: CO₂, H₂O, O₂, N₂, SO₂, Ar;
• Non-CO₂ gases had the same probability of getting dragged into the CO₂-rich stream unintentionally.

2.5. Methanol Synthesis

The methanol synthesis is the facility’s core operation unit. Model input parameters were process streams given by the CO₂ processing unit (DeOxo reactor or CO₂ capture), as well as H₂ generated by the alkaline electrolyzer. Methanol reaching a purity of 99.85 wt% (dry basis), as stated by the IMPCA [38], is obtained at the distillation column’s head. Water, a byproduct of CO₂ hydrogenation, is obtained at the column’s bottom and recirculated to the electrolyzer surpassing a water treatment unit. Significant parameters of the methanol synthesis model were:

• CO₂ conversion rate ψ = 98.11% [39];
• H₂:CO₂ ratio ω > 3.

Detailed information regarding the design and modeling of methanol synthesis units valorizing CO₂ can be found in [3,8,9,17,40,41].

Methanol is synthesized in a single multi-tubular fixed-bed BWR with closed-loop configuration. The evaporation of boiling water ensures an isothermal operation by transferring reaction heat. Catalyst bed cooling is mandatory, since agglomeration of catalyst particles (a conventional Cu/ZnO/Al₂O₃ catalyst is applied for methanol synthesis) occurs as soon as a critical temperature level inside the reactor is exceeded [40]. The reactor’s input streams are CO₂, H₂ as well as unconverted gases (a mixture of CO₂, CO and H₂). The output stream consists of methanol, water and unconverted gases being conveyed to a downstream flash drum. The main chemical reactions occurring are the following:

$${\mathrm{CO}}_2+3{\mathrm{H}}_2\leftrightarrow {\mathrm{CH}}_3\mathrm{OH}+{\mathrm{H}}_2\mathrm{O} \Delta H_r^{298K,50bar}=-40.9\frac{kJ}{mol}.$$

(11)

$${\mathrm{CO}}_2+{ \mathrm{H}}_2\leftrightarrow \mathrm{CO}+{\mathrm{H}}_2\mathrm{O} \Delta H_r^{298K,50bar}=42.0\frac{kJ}{mol}.$$

(12)

Low temperature and high pressure levels favor the conversion of CO₂ following the principle of Le Chatelier. Regarding the temperature inside the reactor, a trade-off between thermodynamics (favored by low temperatures) and kinetics (favored by high temperatures) needs to be found. Furthermore, a critical temperature level must not be exceeded to avoid catalyst deactivation. Preventing the occurrence of the reverse water–gas shift reaction (12) is another key aspect for reactor design, since CO₂ and H₂ are reactants for the main reaction (11). A minimum H₂:CO₂ ratio of 3 must be ensured inside the reactor to prevent the formation of elemental C due to the Boudouard reaction [40]. The reactor was chosen to operate at a temperature of T_MeOH = 513 K and a pressure of p_MeOH = 40 bar.

Separating crude methanol from non-condensable gases is realized by a flash drum following the BWR. Product gas leaving the reactor is cooled to a temperature of T_FD = 323 K by depressurization over a relief valve. Condensed H₂O and methanol are pumped to a downstream distillation column, whereas non converted gases are recycled back to the reactor inlet including a compression step to p_MeOH = 40 bar. The mass flow rate of recirculated gases was calculated with a mass balance over the methanol reactor at steady state (13). The mass flow rate of excess H₂ was calculated by subtracting the required stoichiometric volume flow rate of H₂ from the real volume flow rate of H₂ going into the reactor (14), whereas the mass flow rate of unconverted CO₂ was defined by the CO₂ conversion rate ψ (15). Conversions regarding mass, volume and material balances were conducted with parameters listed in the Appendix A (

Table A4. Molar masses of relevant species.

Species	Molar Mass [kg/kmol]
C	12.01
H2	2.015
O2	32.00
S	32.07
N2	28.01
H2O	18.02
SO2	64.07

).

$${\dot{m}}_{Unc.}={\dot{m}}_{C{O}_2,Unc.}+{\dot{m}}_{H_2,Excess}.$$

(13)

$${\dot{V}}_{H_2,Excess}={\dot{V}}_{H_2,MeOH}-3·{\dot{V}}_{C{O}_2,MeOH}.$$

(14)

$${\dot{n}}_{C{O}_2,Unc.}=\left(1-\psi \right)·{\dot{n}}_{C{O}_2}.$$

(15)

Product separation is realized with a conventional distillation column including the following assumptions:

• T_Feed = 353 K;
• T_Head = 337 K;
• T_Bottom = 373 K;
• Mass fraction of methanol in the feed m’_Feed = 64 wt% (50 mol% methanol);
• Mass fraction of methanol in the head m’_Head = 99.85 wt% [38].

2.6. Alkaline Electrolyzer

The required mass flow rate of H₂ depends on the amount of utilized CO₂ defined by the allocation coefficient γ, the DeOxo reactor’s H₂ demand for process routes including oxyfuel combustion, as well as the chosen H₂:CO₂ ratio ω. The electrolyzer’s power and water demands were defined by the mass flow rate of valorized CO₂. In addition to H₂, O₂ is generated by water splitting, which can either be used as oxidant for process routes, including oxyfuel combustion, or be sold as by-product.

The underlying chemical reactions of water splitting for alkaline electrolysis cells, as well as further information regarding electrolytes, can be found in [42].

Conventional alkaline electrolyzers operate at pressure levels ranging from 1 to 30 bar and at temperatures below 353 K to prevent material degradation [43].

The following assumptions were made:

• p_Electrolyzer = 32 bar;
• T_Electrolyzer = 353 K;
• Specific power demand of 4.45 kWh/Nm³ H₂;
• Specific water demand of 0.804 kg/Nm³ H₂.

2.7. Energy Balance Model

2.7.1. Global Efficiency and Power-to-Liquid Efficiency

The plant’s global efficiency rate η_Tot. includes chemical energy stored in produced methanol ${\dot{U}}_{MeOH}$ and heat provided to adjacent villages via a district heating network ${\dot{Q}}_{DH}$ (Outputs), as well as chemical energy stored in woodchips ${\dot{U}}_{\mathit{Woodchips}}$, the electrolyzer’s electricity demand P_Electrolyzer and the ASU power input P_ASU (Inputs) (16). Chemical energy stored in woodchips and methanol was calculated with Equation (17).

$${\eta }_{Tot.}=\frac{{\dot{U}}_{MeOH}+{\dot{Q}}_{DH}}{{\dot{U}}_{Woodchips}+P_{Electrolyzer}+P_{ASU}}.$$

(16)

$${\dot{U}}_{\mathit{Fuel}/\mathit{Product}} = {\dot{m}}_{\mathit{Fuel}/\mathit{Product}}·{\mathit{LHV}}_{\mathit{Fuel}/\mathit{Product}}$$

(17)

The plant’s power-to-liquid efficiency is defined by Equation (18) [44].

$$PtL-efficiency=\frac{Fuel output}{Electricity input}.$$

(18)

2.7.2. Combined Heat and Power (CHP) Process

Since the plant needs to be able to cover its on-site demand, even if no wind is available, adding a steam turbine for power generation is mandatory. Generated electricity surpassing the plant’s on-site demand is used to power the alkaline electrolyzer. CHP output parameters undergo seasonal variations. In winter months the focus is to provide as much heat as possible for nearby villages, whereas the power output will be maximized in summer. The following maximum efficiencies, related to the thermal input into the BFBC, were assumed:

• Maximum electric efficiency η_CHP,el.,max. = 25%;
• Maximum thermal efficiency η_CHP,th.,max. = 65%.

2.7.3. Wind Park

Wind power should be prioritized over solar power due to the lower specific CO₂-equivalent per kg H₂ generated [40]. Manufacturers of commercial wind turbines generally list their products’ performance as rated power, and add graphs displaying the turbine’s annual energy production as the function of the average wind speed at hub height. Calculating real annual average power outputs of wind turbines can be done for annual average wind velocities of certain regions. A hyperlink to the Austrian Windatlas can be found in [45], indicating annual average wind velocities at a height of 100 m above ground level.

The following assumptions, regarding a wind park adjacent to the designed power-to-green methanol plant, were made:

• Annual average wind velocity u_Wind = 7.5 m/s;
• 12 wind turbines with a total rated power of P_{Wind park} = 36 MWp;
• Annual average power level of P_{Wind park} = 17.14 MW at u_Wind = 7.5 m/s (47.6% of the rated power).

2.7.4. Fluidization of Fuel and Bed Material

To ensure the fluidization of fuel and bed material, a blower is required for overcoming the pressure difference caused by the bed itself, the distributor plate, a baghouse filter, a SCR-DeNO_x unit and a condenser, as indicated in Equation (19). The pressure drop due to the fluidization of fuel and bed material was calculated with Equation (20). Typical properties of silica sand as bed material are listed in the Appendix A (

Table A8. Assumed properties of silica sand as bed material.

Parameter	Value [Unit]
ρP	2660 kg/m3
ρBulk	1490 kg/m3
ε	0.44
dP	0.7 mm
dSV	0.61 mm

).

$$\Delta p_{Blower}=\Delta p_{Bed}+\Delta p_{Dis.}+\Delta p_{Bag.}+\Delta p_{SCR}+\Delta p_{Con.}$$

(19)

$$\Delta p_{Bed}=\left(1-\epsilon \right)·\left({\rho }_P-{\rho }_g\right)·g·h_{Bed}.$$

(20)

The following assumptions were made:

• 80% of Δp_BFBC were caused by fuel and bed;
• 20% of Δp_BFBC were caused by the distributor plate;
• Pressure drop caused by the baghouse filter Δp_Bag. = 2000 Pa;
• Pressure drop caused by the SCR-DeNO_x unit Δp_SCR = 500 Pa;
• Pressure drop caused by the condenser Δp_Con. = 5000 Pa;
• Blower efficiency η_Blower = 0.9.

Consequently, the required power could be calculated using Equation (21). A power reserve was added to ensure smooth operation for real applications.

$$P_{Blower}=\frac{\dot{V}·\Delta p_{Blower}}{{\eta }_{Blower}}+P_{Reserve}.$$

(21)

2.7.5. Conveyor Units

A screw conveyor, fans and pumps were added to the model to estimate the plant’s on-site demand. Compared to major energy consuming units (i.e., compressors, alkaline electrolyzer and ASU), the power demand of conveyor units was almost negligible. Estimations, using conventional models to calculate parameters like the pipe friction factor ξ, the Reynolds number Re and the pressure drop per length unit Δp, were added. The screw conveyor’s power demand was set constant at a value of 0.7 kW for the assumed input of woodchips, based on an online tool calculating the power demand of bulk handling systems [46].

2.7.6. Compressors

Compressors are needed to realize required pressure levels inside the DeOxo (p_DeOxo = 10 bar) and methanol synthesis reactor (p_MeOH = 40 bar). The following assumptions were made:

• Gases behave like ideal gases;
• Isothermal change of states;
• Compressor efficiency η_Compr. = 0.9;
• The stream of unconverted, recirculated gases consists of 33.33 vol% of H₂, CO and CO₂, respectively.

The specific technical work for compressing gas streams was calculated by Equation (22). Consequently, the required power could be calculated by Equation (23). Values regarding specific gas constants of relevant species R_i can be found in the Appendix A (

Table A5. Specific and universal gas constants.

Species	Gas Constant [J/(mol∙K)]
RUniv.	8.314
CO2	188.91
H2	4126.05
CO	296.82
Rec. Gases (H2, CO, CO2)	1084.06

).

$$w_t=R_i·T·\ln \left(\frac{p_2}{p_1}\right).$$

(22)

$$P_{Compr.}=\frac{{\dot{m}}_i·w_t}{{\eta }_{Compr.}}.$$

(23)

2.7.7. Air Separation Unit

For the implementation of process routes involving oxyfuel combustion, the addition of an ASU is mandatory, since plant operation must be independent from daily and seasonal wind fluctuations. If u_Wind falls below a threshold of 3–4 m/s, no power is generated by the wind park and hence no O₂ is generated by the electrolyzer. As a result, the ASU should be connected to the local power grid to avoid plant downtime. For simplification purposes, only N₂, O₂ and Ar output streams were considered. Additionally, a specific energy consumption of 0.45 kWh/Nm³ O₂ was assumed. Finding the required volume flow stream of O₂ provided by the ASU was done with Equation (24).

$${\dot{V}}_{ASU,O_2}={\dot{V}}_{Ox.,O_2}-{\dot{V}}_{Electrolyzer,O_2}.$$

(24)

3. Results and Discussion

After modeling subprocesses and linking them together, several operating points were defined by varying critical model parameters (mainly the current wind velocity u_wind and, consequently, the possible share of utilized CO₂ γ). Mass, material, volume and energy balances were evaluated for defined operating points. Values assumed for the CHP process refer to an operation during winter months, aiming to maximize the output of heat transferred to a district heating network (${\dot{Q}}_{CHP}$ = 13 MW_th., P_CHP = 3 MW). In this work the total amount of electricity generated by an adjacent wind park (rated power of P_{Wind park} = 36 MWp) and power generated by the CHP process was used for H₂ production. Using only surplus electricity, not required by supplied villages, might pose a promising alternative process configuration. However, this case was not considered in this paper.

3.1. Oxyfuel Combustion

Parameters regarding oxyfuel combustion (α, β, r, T_Con.) are constant for all operating points due to a constant thermal input of 20 MW_th.. Figure 2 shows a detailed flow chart of the oxyfuel combustion model including mass and volume balances. A total input of 6133 kg/h woodchips (thermal load of 20 MW_th.) with properties assumed in Section 2.1 and 6187 kg/h O₂ is required. A recirculation ratio of r = 4 ensures a fluidization velocity of u_f = 1–2 m/s (21,664 Nm³/h of fluidization medium). Since biomass requires a relatively high amount of O₂ in the fluidization medium to avoid incomplete combustion, a value of α = 1.08 was chosen (β = 1.4). The volume flow rate of generated flue gas is ${\dot{V}}_{FG}$ = 4333 Nm³/h, obtaining a CO₂ concentration of 89.8 vol% at a temperature of 288 K. Main impurities are excess O₂ (7.41 vol%), H₂O (2.48 vol%) and N₂ (0.31 vol%). A mass flow rate of 4066 kg/h condensed water is pumped to the alkaline electrolyzer after passing a water treatment unit. A total of 2.54 kg/h NH₄OH and 7.23 kg/h Ca(OH)₂ are required for the removal of NO_x and SO_x. A baghouse filter was added to the flowchart for dust removal without being analyzed in further detail.

3.2. Reference Operating Point—Oxyfuel Combustion

The “reference operating point” serves as a benchmark for process routes, including oxyfuel combustion, with all parameters set at average values. As aforementioned, the wind park consists of 12 windmills with a rated power of 3 MWp each. Parameter u_Wind was assumed to be 7.5 m/s at hub height corresponding to the Austrian Windatlas [45]. A total of 1637.4 kg/h methanol is produced utilizing 30% (about 2300 kg/h) of generated CO₂ (γ = 0.3). Process indicators, including specific consumptions per kg methanol produced, can be found in

Table 3. Process indicators regarding the production of methanol—reference operating point.

Process Indicators	kg/h	kg/kg Fuel	kWhel. Consumed/kg	kg/kg H2	kg/kg O2	kg/kg CO2
Methanol	1637.4	0.267	10.76	4.792	0.265	0.306

. A total of 10.76 kWh of power is required per kg methanol produced. Key parameters for the reference operating point are listed in

Table 4. Key parameters—reference operating point.

Parameter	Value [Unit]
Thermal input	20 MWth.
Woodchips input	6133 kg/h
Water content	30 wt%
α	1.08
β	1.40
FGR	Dry
r	4.0
TCon.	15 °C
γ	0.3
CO2 conversion ψ	98.11%
PWind park	36 MWp
PElectrolyzer	17.14 MW
uWind	7.5 m/s

. To meet the BFBC’s total O₂ demand of 6187 kg/h, 3437 kg/h O₂ (55.55%) needs to be provided by an ASU (in addition to 2713 kg/h O₂ (44.45%) provided by the electrolyzer), resulting in a power demand of P_ASU = 4.73 MW. The facility’s input and output streams can be found in the Appendix A (

Table A2. Input and output—reference operating point.

	Parameter	Value [kg/h]	kg/kg Fuel	kg/kg MeOH
Input	${\dot{m}}_{Fuel}$	6132.90	-	3.746
	${\dot{m}}_{H_2}$	341.68	0.056	0.209
	MeOH synthesis (H2)	324.36	0.053	0.198
	DeOxo reactor (H2)	17.31	0.003	0.011
	${\dot{m}}_{O_2}$ (Oxyfuel)	6186.95	1.009	3.779
	O2 from Electrolyzer	2713.07	0.442	1.657
	O2 from ASU	3473.83	0.556	2.122
	${\dot{m}}_{Ca{\left(OH\right)}_2}$	7.226	0.001	0.004
	${\dot{m}}_{N{H}_{4}OH}$	2.447	0.0004	0.001
	${\dot{m}}_{Air}$	14,999	2.446	9.160
Output	${\dot{m}}_{MeOH}$	1637.40	0.267	-
	${\dot{m}}_{N_2}$	11,334.45	1.848	6.992
	${\dot{m}}_{Ar}$	190.33	0.031	0.116
	${\dot{m}}_{FG}$	5744.46	0.937	3.508
	${\dot{m}}_{C{O}_2}$	5348.72	0.872	3.267
	${\dot{m}}_{H_{2}O}$ (net value)	1929.95	0.315	1.179
CO2	CO2 generated	7461.03	1.246	4.667
	CO2 utilized	2292.31	0.374	1.400
	CO2 emitted	5348.72	0.872	3.267

). The on-site demand of 0.72 MW is calculated in

Table A3. On-site demand—reference operating point.

	Process	Symbol	Type	Species	[kg/s]	[Nm3/s]	Power [kW]
Oxyfuel	OXY.1	V1	Blower	FM	10.83	6.02	500
	OXY.2	V2	Fan	FG	0.68	0.42	0.75
	OXY.3	V3	Fan	FG	1.60	0.84	1.5
	OXY.4	P1	Pump	H2O	1.13	-	0.4
	OXY.5	H1	Twin screw	Fuel	1.70	-	0.7
					SUM Oxyfuel		503.35
DeOxo	DeOxo.1	V2	Fan	FG	See OXY.2
	DeOxo.2	C4	Compressor	FG	0.684	-	95.24
					SUM DeOxo		95.24
Methanol	MeOH.1	P2	Pump	MeOH + H2O	0.71	-	0.4
	MeOH.2	P3	Pump	MeOH	0.46	-	0.4
	MeOH.3	P4	Pump	H2O	0.26	-	0.4
	MeOH.4	P5	Pump	H2O	0.26	-	0.4
	MeOH.5	C1	Compressor	CO2	0.64	-	59.87
	MeOH.6	C2	Compressor	H2	0.09	-	32.55
	MeOH.7	C3	Compressor	Rec. gases	0.019	-	27.94
					SUM Methanol		121.96
Electrolyzer	ELE.1	P1	Pump	H2O	See OXY.4
					SUM Electrolyzer		0
				Total on-site power demand			720.55

, listed in the Appendix A. A global power balance is listed in

Table 5. Global power balance.

Name	Power [MW]
Power generation (CHP)	3.0
On-site demand	−0.72
PElectrolyzer	−17.14
PASU	−4.73

, showing a power input into the electrolyzer of P_Electrolyzer = 17.14 MW, as well as a power input into the ASU of P_ASU = 4.73 MW.

Results regarding the subprocesses are visualized in the figures below. The mass and material balance of the DeOxo model, using parameters stated in Table 4, are illustrated in Figure 3. For γ = 0.3, a total of ${\dot{V}}_{FG}$ = 1300 Nm³/h with a CO₂ concentration of 𝑦_CO2 = 89.8 vol% needs to be processed by the DeOxo reactor. Excess O₂, with a material flow rate of ${\dot{n}}_{O_2}$= 4.3 kmol/h (7.41%), is separated catalytically with H₂ generated by the alkaline electrolyzer (${\dot{V}}_{H_2}$ = 193 Nm³/h). A total amount of 2292 kg/h CO₂ with a purity of 99.66 vol% is provided for the downstream methanol synthesis process step.

A material and mass balance of the methanol synthesis for the reference operating point is shown in Figure 4. A total of 52.09 kmol/h of CO₂, provided by the upstream DeOxo reactor, as well as recirculated unconverted gases are transferred into the reactor alongside 160.97 kmol/h H₂ (resulting in a H₂:CO₂ ratio of ω = 3.09). The gas mixture containing CO₂, CO, H₂ and small impurities of N₂ is compressed before passing the Cu/ZnO/Al₂O₃ catalyst bed inside the BWR’s tubes, at a temperature of T_MeOH = 240 °C and a pressure of p_MeOH = 40 bar. The reaction products, methanol and H₂O, are then separated from non-condensable gases in a flash drum by depressurization to p_Atm.. Subsequently, crude methanol is separated into methanol (1637 kg/h) and H₂O (920 kg/h) in a multi-staged distillation column.

H₂ required for CO₂ purification in the DeOxo reactor and CO₂ hydrogenation is generated by an alkaline electrolyzer. A total of 3056 kg/h of H₂O and 17.14 MW power are required to produce ${\dot{V}}_{H_2}$= 3800 Nm³/h and ${\dot{V}}_{O_2}$= 1900 Nm³/h at process conditions of T_Electrolyzer = 80 °C and p_Electrolyzer = 32 bar. Figure 5 shows the alkaline electrolyzer’s balance for the reference operating point.

3.3. Air Combustion and Methanol Synthesis—Operating Point “Air”

Determining whether oxyfuel combustion or conventional air combustion should be prioritized to maximize the facility’s efficiency rates is a key aspect of this paper. The air combustion model is based on the oxyfuel model with some adaptions (i.e., using λ as the combustion parameter instead of α and applying a MEA absorption process for CO₂ processing).

Table 6. Process indicators regarding the production of methanol—operating point “Air”.

Process Indicators	kg/h	kg/kg Fuel	kWhel. Consumed/kg	kg/kg H2	kg/kg CO2
Methanol	1719.23	0.280	9.829	5.032	0.328

and

Table 7. Key parameters—operating point “Air”.

Parameter	Value [Unit]
Thermal input	20 MWth.
Woodchips input	6133 kg/h
Water content	30 wt%
λ	1.2
TCon.	15 °C
γ	0.35
ηCapture	0.9
Purity CO2 stream	99 vol%
CO2 conversion ψ	98.11%
PWind park	36 MWp
PElectrolyzer	17.14 MW
uWind	7.5 m/s

present process indicators as well as key parameters regarding the “Air” operating point. Compared to the reference operating point applying oxyfuel combustion, the specific power consumption of 9.829 kWh_el./kg methanol is significantly lower since no additional H₂ is required for flue gas processing. A global power balance is stated in

Table 8. Global power balance—operating point “Air”.

Name	Power [MW]
Power generation (CHP)	3.0
On-site demand	−0.80
PElectrolyzer	−17.14

. The plant’s on-site demand is calculated similarly, as explained previously for the “reference operating point”. The facility’s on-site demand of 0.8 MW is slightly higher compared to the reference operating point’s on-site demand of 0.72 MW, since more power is required for gas compression due to a higher CO₂ utilization rate of γ = 0.35. This configuration obtains a methanol output of about ${\dot{m}}_{MeOH}$ = 1719 kg/h.

Mass and volume balances of the air combustion model including flue gas treatment units can be found in Figure 6. Since air is used as fluidization medium, the amount and composition of emitted flue gas differ significantly in comparison with oxyfuel combustion. A volume flow rate of ${\dot{V}}_{FG}$ = 27,175 Nm³/h with a CO₂ concentration of only 14.33 vol% is either emitted to the atmosphere or processed in a downstream CO₂ capture unit. 22,968 Nm³/h air are required to achieve an air ratio of λ = 1.2. A total of 4066 kg/h of condensed water is pumped to the electrolyzer surpassing a water treatment unit. The balance of MEA absorption cycle processing of 35% (γ = 0.35) of generated flue gas can be found in Figure 7, processing a volume flow rate of ${\dot{V}}_{FG}$ = 9511 Nm³/h in the absorber at T_Abs. = 308 K. CO₂ is absorbed by a cycling MEA solution (${\dot{m}}_{MEA}$ = 53,136 kg/h) and subsequently released in the desorber at T_Des. = 393 K, demanding 2.54 MW thermal power for regeneration. A mass flow rate of purified CO₂ ${\dot{m}}_{C{O}_2}$= 2407 kg/h with a purity larger than 99 vol% is conveyed to the downstream methanol synthesis process step. Since small shares of CO₂ and other components are not absorbed, ${\dot{V}}_{FG,Lean}$ = 8137 Nm³/h is emitted to the atmosphere with a CO₂ concentration of 2.54 vol%.

After being purified, the CO₂ stream is valorized to methanol, as explained in Section 2.5. The reactor’s input streams are ${\dot{n}}_{C{O}_2}$= 54.69 kmol/h and ${\dot{n}}_{H_2}$= 169.57 kmol/h, resulting in a H₂:CO₂ ratio of ω = 3.1. After being separated in a distillation column, a product stream of 1719 kg/h methanol and 966 kg/h H₂O are acquired. Figure 8 shows the methanol synthesis’ balance.

Similar to the oxyfuel process route, H₂ is provided by an alkaline electrolyzer powered by a wind park. Input parameters (${\dot{m}}_{H_{2}O}$= 3056 kg/h and P_Electrolyzer = 17.14 MW) and output parameters (${\dot{V}}_{H_2}$= 3800 Nm³/h and ${\dot{V}}_{O_2}$= 1900 Nm³/h) are similar to the reference operating point described in Section 3.2.

3.4. Sensitivity Analysis

A sensitivity analysis was conducted to present how critical model parameters (e.g., α, γ, ω, u_Wind) affect the facility’s performance.

3.4.1. O₂ Supply for Oxyfuel Process Routes

For process routes including oxyfuel combustion, a constant mass flow rate of ${\dot{m}}_{O_2}$ = 6187 kg/h must be provided to maintain the combustion reaction (α = 1.08, β = 1.4, ${\dot{m}}_{Woodchips}$ = 6133 kg/h, thermal heat load = 20 MW_th.,). The mass flow rate of O₂ into the combustor, as well as the flue gas’ O₂ concentration, are a function of α. Higher O₂ concentrations lead to an increased H₂ demand of the DeOxo reactor and hence results in the plant’s performance parameters’ decrease. Figure 9 shows how ${\dot{m}}_{O_2}$ (blue line) and $y_{FG, O_2}$ (red line) are influenced by α. Small values of α conclude in a smaller power demand of ASU and CO₂ purification; however, they will possibly lead to incomplete combustion regimes.

The provided volume flow rate of O₂ is a mixture of O₂ generated by an ASU and the electrolyzer as by-product of water splitting. Figure 10 displays the share of O₂ provided for the oxyfuel combustion by an ASU and the alkaline electrolyzer, with a power input of P_Electrolyzer = 17.14 MW, for the reference operating point (γ = 0.3, see Section 3.2). An O₂ stream of 1925.84 Nm³/h is provided by the electrolyzer (44.45%), 2406.96 Nm³/h is provided by the ASU (55.55%) to meet the oxyfuel combustor’s total demand of 4332.8 Nm³/h as indicated by the green line. The “break-even point” of O₂ supply is reached at P_Electrolyzer = 38.56 MW. Consequently, oxyfuel process routes do not require an additional ASU if γ > 0.7, hence operating points surpassing an allocation coefficient of 0.7 produce excess O₂ as by-product, making oxyfuel combustion a viable option for large-scale applications.

3.4.2. Methanol Production and Analyzed Operating Points

H₂ production was identified as the facility’s bottleneck defining the output of methanol. The amount of H₂ provided by the electrolyzer depends on the electrolyzer’s size (=const.), the wind park’s size (=const.) and the wind velocity (≠const.). Depending on these parameters, the allocation coefficient γ can reach values between 0 (no flue gas is processed to methanol) and 1 (the total amount of flue gas is processed to methanol). Besides operating points explained previously (reference operating point and operating point “Air”), additional operating points were analyzed by varying u_Wind, γ and P_{Wind park}, as displayed in

Table 9. Analyzed operating points.

Operating Point	PWind park [MWp]	uWind [m/s]	γ [-]	${\dot{\mathit{m}}}_{\mathit{M}\mathit{e}\mathit{O}\mathit{H}} [\mathbf{kg}/\mathbf{h}]$
Reference operating point	36	7.5	0.3	1637.4
Max. output of methanol “Max. MeOH”	123	7.5	1	5457.9
Low wind velocity “Min Wind”	36	4	0.1	545.8
Max. wind velocity “Max. Wind”	36	12	0.6	3274.7
Air combustion mode “Air”	36	7.5	0.35	1719.2
Biomass heating plant “Dead Calm”	36	0	0.04	196.5

. Operating point “Dead Calm” refers to a scenario with no power generated by the wind park due to u_Wind = 0 m/s, thus the electrolyzer is only powered with surplus electricity generated by the CHP process (P_Electrolyzer = 2 MW). The maximum yield of methanol can be produced in combination with a wind park at a rated power of P_{Wind park} = 123 MWp and u_Wind = 7.5 m/s, as indicated by operating point “Max.MeOH”. Processing the whole amount of generated CO₂ (γ = 1.0) would require a power input into the electrolyzer of about 60 MW. Therefore, this scenario is not realistic regarding the size of currently used electrolyzer units. At u_Wind = 4 m/s, the lowest possible wind velocity for wind turbines to operate economically feasible, a yield of 545.8 kg/h of methanol can be produced (γ = 0.1, P_{Wind park} = 36 MWp). A total of 3274.7 kg/h of methanol can be produced at wind velocities surpassing u_Wind = 12 m/s (γ = 0.6, P_{Wind park} = 36 MWp).

The mass flow rate of produced methanol (green line), as well as the required power input of the electrolyzer (red line) as a function of allocation coefficient γ, can be found in Figure 11.

3.4.3. H₂ Demand

Producing renewable H₂ is the determining factor when regarding the conceptualized power-to-green methanol plant’s power consumption. Stoichiometrically, a H₂:CO₂ ratio ω larger than 3 is mandatory. For real applications, ω is expected to be set at values ranging from 3 to 3.5. Increasing mass flow rates of H₂ for methanol production as a function of ω are displayed in Figure 12. The electrolyzer’s required power input ranges from P_Electrolyzer = 15.58 MW (ω = 3, ${\dot{m}}_{H_2}$ = 314.86 kg/h) to P_Electrolyzer = 18.18 MW (ω = 3.5, ${\dot{m}}_{H_2}$ = 367.34 kg/h) for the reference operating point, as stated in Section 3.2 (γ = 0.3). The reader should keep in mind that these values only serve as benchmarks, since recirculated gases (H₂, CO, and CO), which are not discussed in this concept, will have a significant impact on the reactor’s operating parameters for real application scenarios.

3.5. Global Efficiency and Power-to-Liquid Efficiency

Obtained efficiency rates for operating points presented in Section 3.4.2 are listed in

Table 10. Efficiency rates for analyzed operating points (As defined in Section 2.7.1).

Operating Point	ηTot. [%]	ηPtL [%]	γ [-]	${\dot{\mathit{m}}}_{\mathit{M}\mathit{e}\mathit{O}\mathit{H}} [\mathrm{kg}/\mathrm{h}]$
Low wind velocity “Min Wind”	47.28	20.97	0.1	545.8
Reference operating point	53.19	40.19	0.3	1637.4
Max. wind velocity “Max. Wind”	57.04	51.10	0.6	3274.7
Max. output of methanol “Max. MeOH”	56.09	51.91	1	5457.9
Air combustion mode “Air”	54.36	54.21	0.35	1719.2

. When analyzing process routes including oxyfuel combustion, η_Tot.,Oxy increases with increasing γ, obtaining a maximum of η_Tot.,Oxy = 57.04% (γ = 0.6) for the “Max. Wind” operating point. η_PtL,Oxy increases with increasing γ, obtaining a maximum of η_PtL,Oxy = 51.91% (γ = 1) for the operating point “Max MeOH”. Increasing the wind park’s size (γ↑); therefore, leads to increased efficiency rates, η_Tot.,Oxy and η_PtL,Oxy, for oxyfuel process routes. This can be explained by the fact that no additional O₂ needs to be provided by an ASU if γ > 0.7 (P_Electrolyzer > 38.56 MW), as mentioned in Section 3.4.1, resulting in a significantly decreasing global power consumption. Therefore, power-to-green methanol plants based on oxyfuel combustion routes obtain significantly higher efficiency rates at a large scale.

Significant deviations can be found when comparing process routes, including oxyfuel combustion (reference operating point) to conventional air combustion (“Air”), obtaining comparable methanol product output streams (${\dot{m}}_{Reference}$ = 1637.4 kg/h, ${\dot{m}}_{Air}$ = 1719.2 kg/h). The PtL efficiency of operating point “Air” η_PtL,Air = 54.21% is substantially higher than the PtL efficiency of the reference operating point η_{PtL,Reference} = 40.91%. Two reasons need to be highlighted: Firstly, the ASU’s high-power consumption, providing 44.45% of the combustion’s O₂ demand, has a negative impact on η_PtL. Secondly, H₂ needs to be provided for CO₂ processing in the DeOxo reactor, resulting in a significant penalty on η_PtL.

3.6. Case Study for One Year of Operation

The power-to-green methanol plant’s expected annual input and output parameters are determined in this Section. Adjacent villages are supplied with heat generated by the CHP process via a district heating network. Electricity generated by the CHP process and the wind park is used for the generation of green H₂ needed to produce methanol. It was assumed that electricity is solely used to power the alkaline electrolyzer. Alternatively, surplus electricity generated by the wind park may be stored chemically by producing green methanol. The following assumptions were made:

• The CHP process output undergoes seasonal fluctuations obtaining a total efficiency of η_CHP,Tot. = 80% with respect to a thermal input of 20 MW_th.;
• Electricity generated by the CHP process is used to power the alkaline electrolyzer;
• The plant’s on-site power demand was set at 0.72 MW, referring to the reference operating point analyzed in Section 3.2;
• The plant is out of operation for three consecutive weeks in June for revision and maintenance work (8256 operating hours per year);
• P_{Wind park} = 17.14 MW (annual average output—reference operating point);
• Seasonal fluctuations of the wind park’s power generation in central Europe were based on Figure A1 in the Appendix A;
• Power consumption of 10.32 kWh/kg methanol produced;
• Power consumption of 4.415 kWh/year and household;
• Heat demand of 1.8 kW/household in winter (central Europe).

CHP operation parameters for one year of operation are listed in

Table 11. CHP process—annual output.

Season	Months	PCHP [MW]	${\dot{\mathit{Q}}}_{\mathit{C}\mathit{H}\mathit{P}} [\mathbf{MW}{\mathbf{th}}_{.}]$	Op. Time [h]	PTot. [MWh]	QTot. [MWhth.]
Winter	Dec, Jan, Feb	2.28	13	2160	4924	28,080
Spring	Mar, Apr, May	3.28	12	2208	7242	26,496
Summer	Jun, Jul, Aug	4.28	11	1704 1	7293	18,744
Autumn	Sep, Oct, Nov	3.28	12	2184	7163	26,208
SUM	-	-	-	8256	26,622	99,528

¹ No operation for three weeks due to revision and maintenance work.

, including an annual heat output of 99,528 MWh_th. and an annual power output of 26,622 MWh. The power output will be maximized in summer since no heat needs to be transferred to a district heating network. The wind park’s monthly power generation over the year is listed in

Table 12. Wind park—annual output.

Month	Wind Index [-]	PWind park [MW]	Op. Time [h]	PTot. [MWh]
Jan	149	25.53	744	18,894
Feb	128	21.93	672	14,737
Mar	133	22.79	744	16,956
Apr	93	15.94	720	11,477
May	81	13.88	744	10,327
Jun	68	11.65	216 1	2516
Jul	77	13.19	744	9813
Aug	73	12.51	744	9307
Sep	90	15.42	720	11,102
Oct	107	18.34	744	13,645
Nov	126	21.59	720	15,545
Dec	126	21.59	744	16,063
SUM	-	-	8256	150,382

¹ No operation for three weeks due to revision and maintenance work.

. A total of 150,382 MWh of power was expected to be generated by the wind park per year. The wind index indicates the wind park’s potential to generate power throughout the year for Central Europe, obtaining its maximum throughout winter months. Plant revision and maintenance was scheduled in June considering that the wind park’s power output is expected to be lowest during summer.

Table 13. Annual input and output—one year of operation.

Season	Months	Operating Time [h]	Fuel Input [t]	Methanol Output [t]	Power equ. [GWh]	Heat Output [GWhth.]
Winter	Dec, Jan, Feb	2160	13,247	5302	54.72	28.08
Spring	Mar, Apr, May	2208	13,541	4457	46.00	26.50
Summer	Jun, Jul, Aug	1704 1	10,450	2803	28.93	18.74
Autumn	Sep, Oct, Nov	2184	13,394	4599	47.46	26.21
SUM	-	8256	50,632	17,161	177.11	99.53

¹ No operation for three weeks due to revision and maintenance work.

provides a summary, based on the reference operating point presented in Section 3.2, of the expected annual input and output streams of the designed power-to-green methanol plant. For an annual input of 50,632 t/a woodchips, a total of 17,161 t/a methanol, equivalent to 177.11 GWh/a of power, can be produced. Additionally, 99.53 GWh_th./a heat is provided to adjacent settlements via a district heating network.

4. Conclusions

The main objective of this paper was to conceptualize and model a decentralized power-to-green methanol plant including oxyfuel combustion for CCU applications. A biomass heating plant, either operated as oxyfuel or conventional air combustor, including a CHP process with a thermal input of 20 MW_th., provides adjacent villages with heat and functions as a CO₂ source for a downstream methanol synthesis unit. Using woodchips as fuel comes with several advantages (e.g., the potential to achieve a negative CO₂ balance, empowering local economies and being independent from imported fossil energy carriers). Processing the CO₂ stream is either realized by a DeOxo reactor (for oxyfuel process routes) or a MEA temperature swing absorption cycle (for conventional air combustion process routes). Required H₂ is produced by an alkaline electrolyzer powered with electricity generated by a nearby wind park. O₂ as by-product can either be used as oxidant for the oxyfuel combustion or be sold when applying process routes based on air combustion. Methanol is produced by valorizing CO₂ in a multi-tubular BWR including a closed gas loop configuration.

PtX technologies are on the verge of becoming key technologies within this decade, as they neutralize a major disadvantage of renewable energy carriers (i.e., the challenge of chemically storing surplus electricity due to seasonal and daily fluctuations). Furthermore, CO₂ is captured and valorized to fuels or feedstocks for the chemical industry.

A variety of operating points have been defined by alternating critical process parameters (i.e., u_Wind, the wind park’s size and the allocation coefficient γ, as explained in Section 3.4.2). Input and output streams of defined operating points were analyzed. Figure A2 and Figure A4, listed in the Appendix A, show a detailed as well as a simplified flow chart of the reference operating point’s balance, a process route including oxyfuel combustion and a DeOxo reactor for CO₂ processing, with parameters set to values with the highest probability for real applications. A product output of 1637 kg/h methanol as well as 13 MW_th. heat is realized for an input of 6133 kg/h woodchips, P_Electrolyzer = 17.14 MW and P_ASU = 4.73 MW. A total amount of H₂ of 342 kg/h is required for the methanol synthesis unit (324 kg/h) and the DeOxo reactor (18 kg/h). Furthermore, 6187 kg/h of O₂ is required for the oxyfuel combustion. For a CO₂ utilization rate of γ = 0.3, 3474 kg/h O₂ (55.55%) must be provided by an ASU (P_ASU = 4.73 MW). In this case, only 2713 kg/h O₂ (44.45%) is provided by the electrolyzer at a power input of P_Electrolyzer = 17.14 MW. If P_Electrolyzer exceeds a value of 38.56 MW (γ = 0.7), no additional O₂ would be required by an ASU. Global flow charts of a comparable scenario using air combustion can be found in Figure A3 and Figure A5 in the Appendix A. In this scenario, 35% of generated flue gas (9511 Nm³/h with a CO₂ concentration of 14.33 vol%) can be converted to methanol. A yield of 1719 kg/h of methanol can be obtained with a fuel input of 6133 kg/h woodchips. Additionally, no ASU is required to save 4.73 MW of power. It is noted that condensed water is pumped to the electrolyzer, which is not indicated in the simplified flow charts.

This paper also focuses on answering the question whether process routes including oxyfuel combustion should be preferred over air combustion, since O₂ generated by the electrolyzer could be used as oxidant substituting an ASU. To evaluate this assumption, plant efficiency parameters were defined and analyzed. The plant’s power-to-liquid efficiency increases significantly with an increasing share of utilized CO₂ (increasing γ) for oxyfuel combustion applications; nevertheless, it is constant for air combustion process routes (η_Tot.,Air = 54.36%≠f(γ), η_PtL,Air = 54.21%≠f(γ)). The maximum efficiency rates for oxyfuel combustion process routes were η_{Tot.,Max.,Oxy} = 57.04% (γ = 0.6) and η_PtL,Max.,Oxy = 51.91% (γ = 1.0), whereas the minimum efficiency rates are significantly lower at η_{Tot.,Min.,Oxy} = 47.28% (γ = 0.1) and η_PtL,Min.,Oxy = 20.97% (γ = 0.1). This aspect is linked to the ASU’s high power consumption for operating points with low γ values. For plants utilizing less than 70% of the generated flue gas, air combustion process routes should be preferred, as they obtain a higher η_PtL,Air = 54.21% (γ = 0.35) compared to η_{PtL,Oxy,Reference} = 40.91 (γ = 0.3) for the reference operating point. However, oxyfuel applications have the potential to reach a higher η_Tot. (η_Tot.,Max. = 57.04%) in comparison to air combustion modes, since no additional heat is required for regenerating the MEA solution for CO₂ processing.

The presented concept has the potential to work as a carbon sink if the methanol produced is utilized as feedstock in the chemical industry. Wooden biomass is used as fuel, while providing decentralized communities with heat generated by combusting local wood, hence empowering regional development. In case methanol is used as fuel, the proposed concept can be described as almost carbon neutral, as CO₂ bound by wood is subsequently emitted to the atmosphere. Additionally, a business case for small-to-medium sized villages situated in regions obtaining a high wind potential is generated. For applications using surplus electricity, a process route including conventional air combustion in combination with a CO₂ capture unit is highly recommended, since no additional ASU is required. This is also true for applications including small-scale wind parks. Oxyfuel combustion processes have a high potential in combination with water electrolysis, since generated O₂ can be directly used as oxidant. However, this configuration is only reasonable if P_Electrolyzer > 38.56 MW (for power plants with a thermal input of 20 MW_th.), the threshold when the combustor’s O₂ demand can be fully covered with O₂ generated by the electrolyzer. Substituting O₂ with an ASU always comes with a significant penalty on η_PtL and should consequently be avoided. As a result, oxyfuel process configurations are only recommended for large-scale investments, including wind parks with a rated power of P_{Wind park} > 100 MWp and large annual average u_Wind.

In addition, a case study analyzing the facility’s annual major input and output streams was conducted assuming process parameters of the reference operating point presented in Section 3.2. An annual output of 17,161 t/a methanol and 99.53 GWh_th./a heat can be realized for an input of 50,632 t/a woodchips and 177.11 GWh/a power, generated by a local wind park and the CHP process.

Combining oxyfuel combustion with CCU technologies (e.g., the valorization of CO₂ to methanol) is a promising system configuration for future PtL plants which use large-scale electrolyzers. Scaling up water electrolysis plants is going to be one of the major challenges for engineers in this decade. The combination of solid oxide electrolysis units and CCU technologies has the potential to significantly improve the efficiency rates of PtL plants by thermal process integration. Heat generated by the combustor (oxyfuel or air) and the methanol synthesis unit can be transferred to the SOEC (solid oxide electrolysis cell) unit, which results in the decrease of the required power for water splitting.

Future research should include the following aspects:

• Basic and detail engineering for all unit operations as well as their interaction;
• Implementation of heat integration by, for example, conducting a pinch analysis;
• Finding new process routes (e.g., using photovoltaic modules as power source, using other CO₂ sources like direct air capture, pre combustion or industrial processes). Exchanging the methanol synthesis unit with a Fischer-Tropsch synthesis or methanation unit are possible options if other products are preferred by possible customers. Additionally, process routes including SOEC units have the potential to significantly improve the facility’ efficiency rates;
• Evaluating the facility’s potential for grid stabilization;
• Developing a technology impact analysis as well as an LCA.

Detailed and simplified global flow charts of process routes applying oxyfuel and air combustion can be found in the Appendix A.