1. Introduction
In recent years, global warming has been one of the major issues facing humanity. According to the Paris Agreement, the rise in global temperatures must be limited to 1.5 °C compared with pre-industrial levels [
1]. Decarbonising existing processes and technologies and switching to renewables are key to achieving these goals.
The trend towards increasing the use of renewable sources for electricity generation is evident. Although among renewables, hydropower, at 1.25 TW, represents the largest share of global generation capacity, wind and solar dominate the newly added generation capacity. Together, both technologies contributed 90% of all renewable energy generation capacity added in 2022 [
2].
Even though many new wind and solar power plants have been built, there is a new problem arising with renewables. This is the problem of unpredictable supply and demand [
3]. During windy or sunny periods, the supply of electricity is high. In cases where demand for energy is low, this can lead to extremely low, or even negative levelised electricity prices on wholesale markets [
4].
Excess renewable energy can be stored to solve the problem of the oversupply of electricity [
1]. One storage solution is the power-to-gas (P2G) process, where excess electricity is used to produce hydrogen (H
2) via water electrolysis. The hydrogen can then be stored, or injected into natural gas pipelines [
3].
Hydrogen storage poses several challenges. Because it is the smallest molecule, it diffuses relatively quickly through the storage tanks [
5]. It has a very low boiling point (−252.8 °C), which requires cryogenic temperatures for liquefaction [
6], and it also has a very low volumetric energy density. According to the ideal gas law, to supply the same amount of energy as 1 mol of methane, hydrogen must be compressed to 3.318 times the storage pressure of methane [
7]. Storing in a gas state therefore requires pressures from 350 to 700 bar [
6]. The presence of hydrogen in natural gas pipelines also causes the deterioration of the mechanical properties of steel materials, due to a phenomenon called hydrogen embrittlement [
5].
Since direct storage of hydrogen is problematic, storage in the form of methane may be more feasible. With hydrogen, methane can be produced by the Sabatier reaction shown in Equation (1). The produced methane can be referred to as synthetic methane, or synthetic natural gas. The production process can be referred as power-to-methane (P2M) [
1]. Synthetic methane is a substitute for natural gas and can therefore be injected into pipeline networks. It can also be stored in the existing infrastructure for natural gas storage.
In addition to a source of hydrogen, the process requires a source of CO
2. CO
2 sources can be divided into three categories, namely, fossil sources, biogenic sources and ambient air [
8]. Biogenic sources are particularly attractive, as CO
2 already belongs to the natural carbon cycle [
9]. Biogenic sources include anaerobic digesters, bioethanol plants and sewage treatment plants. For P2M plants, anaerobic digesters are a particularly suitable source of CO
2, as they are already widely accepted processes for green gas production. As biogas is CO
2-rich, anaerobic digesters upgrading biogas to pipeline-grade biomethane can serve as a source of otherwise unused CO
2 [
8]. Another promising source is the chemical industry, as many processes provide highly concentrated CO
2 streams, resulting in low capture costs [
1].
Another advantage of biogas as a source of CO
2 is that the whole gas can be used directly in the P2M process. The present CO
2 is consumed in the Sabatier reaction, while CH
4 and water (H
2O) are products of the same reaction [
10]. Potential limitations of the direct methanation may arise when there is a presence of gas contaminants such as siloxanes, hydrogen sulphide (H
2S), carbon monoxide (CO) and ammonia (NH
3) [
11]. The presence of H
2S is particularly problematic, as it causes the poisoning of nickel catalysts [
12], which, due to their low price, are used widely for CO and CO
2 methanation [
13].
The P2M process has already been studied widely. Mostly, solar and wind sources are considered as the electricity provider [
1,
14]. Tripodi et al. developed an integrated, five-stage CO
2 methanation process [
15], Calbry-Muzyka and Schildhauer reviewed the direct methanation of biogas [
10], while Hillestad simulated the direct methanation of biogas [
16]. Catarina Faria et al. investigated the operation of Sabatier reactors in equilibrium [
17], Gandara-Loe et al. performed a water-integrated simulation of the process [
18], and Jürgensen et al. simulated the process, where the reaction heat was used for district heating [
19].
It was observed that negative prices started to appear on the Slovenian energy exchange market SouthPool, which contributed to greater interest in the P2M process. Although much work has already been done, there is not much emphasis on the simultaneous methanation of CO2 and biogas. The P2M process could be used simultaneously for biogas being upgraded to biomethane via direct biogas methanation, while it could also serve for the methanation of pure, otherwise-emitted CO2. The biogas would be provided from a nearby biogas plant, while otherwise-emitted CO2 from various companies would be captured and transported to the site. The produced biogas in biogas plants is usually burned at the site to produce heat and power. For the injection of biogas into the grid, the upgrading processes are needed. Using the biogas as a source of CO2 would eliminate such upgrading processes. It is worth noting that if the aforementioned impurities are present in the stream, they would still need to be removed. In these cases, the biogas stream could be passed through the activated carbon. The process partially eliminates the need for CO2 removal and investment costs for upgrading facilities at the biogas plant. The use of biogas in the methanation also utilises the remaining CO2 in the biogas, which could be emitted to the atmosphere if not captured during the upgrading process. CO2 would be used as the feed for the process, to which the biogas stream from the biogas plant would be added.
The proposed P2M process also includes the modified Rankine cycle for simultaneous power production, which increases the amount of generated electricity compared with a simple version of the cycle, where only one turbine is used. While the literature shows the simulation of the process, not many articles include the operation of the necessary dehydration for the produced synthetic methane. In our study, silica gel was used as a potential adsorbent, since it has low price compared with other adsorbents.
A study was therefore carried out on the possible operation of the P2M process. The main objective was to simulate a process that would take advantage of surplus electricity in the electric grid and electricity from renewable sources, in order to produce synthetic methane at low prices. The proposed process could also operate at negative prices on the energy exchange market. In this case, the CO2 methanation process would be able to produce around 1 t of synthetic methane per hour. The process could produce synthetic methane with pure CO2 methanation, or with simultaneous biogas and CO2 methanation. Two cases were therefore simulated in Aspen Plus. The biogas flow rate was estimated at 26.99 kmol/h based on the biogas production capacity of an industrial plant.
It was considered that the produced synthetic methane could be injected into the pipeline. The product stream was therefore compressed to 51 bar and dehydrated with pressure–temperature swing adsorption (PTSA). For the PTSA, a dynamic simulation was performed in Aspen Adsorption.
For both cases of methanation, the process scheme was identical and was partly heat-integrated. Because the Sabatier reaction is exothermic, the process scheme also involved the simultaneous generation of electricity utilising the released heat.
The produced synthetic methane plays an important role for the energy transition. It can be used as a synthetic fuel, replacing natural gas. Because the two of them are very similar in composition, the existing infrastructure can be used for its storage. The process offers seasonal, or even long-term energy storage, while utilising otherwise-emitted CO2. The process also offers storage for hydrogen, since hydrogen is very difficult to store.
3. Results and Discussion
In the first case, the total CO
2 conversion was 99.43%, while in the second case, it was 99.33%. In the methanation of the first case, higher conversions are achieved due to the Le Chatelier principle, which states that the system responds to the addition of reactants by lowering the concentration of the reactants in equilibrium [
45]. Since CH
4 is a product of the methanation reaction and is already present in the reactant stream at the beginning, the equilibrium effects lead to the formation of a smaller amount of products. In a similar way, the removal of products shifts the equilibrium of the reaction in the direction of the products [
45]. Cooling of the product stream between the reactor stages was implemented for this reason. If one looks at the Sabatier reaction shown in Equation (1), it can be seen that for 1 mol of CH
4 produced, 2 mol of H
2O is produced. The maximum CO
2 conversion is limited by the equilibrium. However, because a lot of water is produced, its condensation from the product stream allows higher CO
2 conversions in the second reaction stage. With a sufficiently large number of reaction steps with intermediate cooling, CO
2 conversions up to even 100% are possible [
15].
Figure 7 shows the conversion–temperature path of the process. It is important to note that lines 2–3 and 5–6, representing reactors R-1 and R-2, respectively, do not show the actual profile of conversion in the reactor, but only the conversion of CO
2 at the inlet and outlet of the reactors. Point 1 represents the temperature of the S-2 stream, the point 4 represents S-6 stream and the point 7 represents S-9 stream. The temperature is higher in case 2, due to the compression of the biogas, which heats the biogas stream in the second compression stage.
3.1. Temperature and Concentration Profiles
The issue of the formation of a hot spot in the first reactor arises. When the simulation was carried out with a 120 kg catalyst, the peak temperatures reached as high as 600 °C. By lowering the mass to 40 kg, the temperature was reduced successfully, but
Figure 8 shows that the hot spot was still present. Lowering the catalyst mass further led to simulation errors.
The formation of the hot spot was caused by a rapid consumption of reactants in the initial lengths of the reactor, as shown in
Figure 9. The hot spot was therefore formed at 1.8 m. At the same location, we observed higher CO fractions. In the adopted model, it is assumed that CO
2 methanation takes place via an intermediate reduction to CO, via the RWGS reaction. This reaction is endothermic, with a standard reaction enthalpy (Δ
rH°) of 41 kJ/mol, while CO methanation is exothermic, with Δ
rH° = −206 kJ/mol [
15]. We concluded that the large amount of heat released during the methanation of CO accelerates the RWGS reaction greatly, while the high temperature makes it more difficult to convert the formed CO further. Because of that, it accumulates in the vicinity of the hot spot. This accumulation can lead to a further reduction of CO to carbon [
20]. This causes a build-up of the active surface of the catalyst, and, with time, leads to a loss of catalytic activity [
46]. In the simulations, we have neglected this, which is why the CO was still converted to CH
4. It would make sense, however, to remove the hot spot and consider possible coke formation in further simulations. To eliminate hot spots, inert material is mixed with the catalyst. Moreover, less catalyst is applied in the initial parts of the reactor [
47]. Aspen Plus, however, does not have a built-in function to add inert material, or to determine a custom distribution of the catalyst mass along the reactor.
In the second reactor, the reactions are also fast, but no hot spot was observed, because much smaller amounts of CO
2 and H
2 were reacting.
Figure 8 also shows the temperature profile of the TF in the reactors.
3.2. Synthetic Methane Dehydration
A six-step cycle was used to simulate 10 dehydration cycles of operation. For each step in the PTSA cycle, the specifications for all valves are presented in
Table 7. In addition, it is shown in which steps of the cycle the column heating was switched on.
Table 8 shows the-time driven PTSA operation. It also shows the operating pressures and the temperature at which regeneration occurred. Final results are shown in
Table 9.
After carrying out the simulation of the first case in Aspen Plus, we obtained the composition of the PROD-2 stream, which is presented next to the other process streams in
Table 10. A dynamic adsorption simulation was performed with the results. Unless otherwise noted, all results in this section refer to the dehydration of the PROD-2 stream, produced with the methanation of pure CO
2. The justification for using the same split fractions for the dehydration of both cases is discussed at the end of this chapter.
The simulated dehydration PTSA ensures a sufficiently low water content in the dehydrated synthetic methane to meet many pipeline standards of different European countries, as given by Awe et al. [
11].
Figure 10 shows how the water fraction and average water fraction in the product stream varies with time. The maximum water fraction in the product stream was 0.00208 mol.%. The determined steady-state value of the average water fraction in the product stream is represented by the horizontal red line. The other steady-state compositions are shown in
Table 9, together with the calculated split fractions.
The steady-state temperatures of the product and waste streams have also been transferred to Aspen Plus. The product stream has warmed up slightly due to the adsorption process in the column. The determined steady-state flows and temperatures were:
= 1.7680 × 10−2 kmol/s;
= 1.7562 × 10−2 kmol/s;
= 1.1848 × 10−4 kmol/s;
= 298.15 K = 25.00 °C;
= 298.97 K = 25.82 °C;
= 371.32 K = 98.17 °C.
According to the split fractions, the methane losses in the PTSA are very low, since 99.42% of all methane from the PROD-2 stream passes into the SYNCH4 stream. According to the second equation, we determined the losses to be 0.57 mol.%. Consequently, the effluent flow rate W was also very low, as can be seen in
Table 10. Such low losses are achieved because only a small fraction of the product stream is passed through the regenerating column. Although this results in long column regeneration times, the chosen column dimensions enable long operating cycles. The column regeneration time for the first cycle is shown by the blue dotted line in
Figure 11.
Via dynamic simulation, we determined that the PTSA dehydration system requires a heat flow of 3.99 kW. The value was obtained by summing the amount of heat exchanged with the heating water for both columns, and then dividing the value by the operating time of the whole system.
For the PROD-2 stream from case 2, cyclic operation was not repeated. Instead, we simulated only the breakthrough curve for the composition and flow of the second case and saw that the use of the dehydration results of the first case was justified.
The solid blue line in the upper diagram in
Figure 11 shows the breakthrough curve for the dehydration in the first case. The flow rate of the PROD-2 stream was higher when dealing with the dehydration of synthetic methane produced by simultaneous methanation. Since the column dimensions were the same in both cases, the breakthrough was therefore achieved more quickly.
To ensure the same maximum water fraction in the product in the dehydration of case 2, the adsorption and desorption times of the columns must be reduced to 56.61 h due to the higher flow rates. From the diagram on the right, approximately the same amount of water would be adsorbed in the column at this time. Since the amount adsorbed is very similar, we assume that this amount would also be desorbed in a very similar time. The time for desorption of such an amount of water can be estimated to 54 h from the blue dotted line on the bottom right diagram. This is sufficiently fast, as the desorption time does not exceed the adsorption time. Therefore, we concluded that the same split fractions can be used for the dehydration of both cases. The expected flow rate and composition of the SYNCH4 stream from case 2 are given in
Table 11.
3.3. Process Stream Results
Table 10 and
Table 11 show the final temperatures, pressures, flows and compositions of all process streams for the simulated synthetic methane production.
Table 10 shows the results of the first case and
Table 11 the results of the second case.
We have simulated the production of synthetic methane successfully. The flow rate of the PROD-2 stream from case 1 was 1013.49 kg/h and corresponds to a calculated pure methane flow rate of 1 t/h. There were minor losses in the PTSA unit, resulting in a final product flow of the SYNCH4 stream of 1006.55 kg/h. In the waste stream of the PTSA unit, 5.75 kg/h of CH4 was lost. Given that the methane flow rate in the PROD-2 stream was 993.85 kg/h, only 0.58% of methane was lost. The value corresponds to the losses predicted by the dynamic PTSA simulation.
The case 1 composition of the SYNCH4 stream corresponds to the literature standards for natural gas pipelines provided by Awe et al. [
11]. The exception is Sweden, because the hydrogen fraction in the final stream is too high. From Aspen Plus, we have also obtained the Wobbe Index on a volume basis at a reference temperature of 0 °C. The Wobbe Index of the SYNCH4 stream is 52.75 MJ/m
3. In some cases, this is too high. For example, the Netherlands requires upper index values between 43.46 and 44.41 MJ/Nm
3. Accordingly, the minimum required CH
4 fraction is also lower, at 85 mol.% [
11]. To lower the Wobbe Index, the synthetic methane product stream could, in such cases, be blended with nitrogen.
Although the CO2 conversions are lower in case 2, the methane fraction of the product is higher. We conclude that this is due to the initial methane content, which raises the methane content of the product. The case 2 compositions of the final stream are within the relevant standards, and the Wobbe index is 52.82 MJ/m3. Due to the initial methane content, the final synthetic methane flow rate is slightly higher, at 1274.53 kg/h. The methane losses in the PTSA are, therefore, also higher (7.30 kg/h), but amount to the same percentage of lost methane (0.58%). It makes sense that the losses are the same, as we have assumed that the split fractions of the PTSA unit in Aspen Plus are the same for both cases.
Methane losses also occur during condensation. Due to dissolution in water, methane is also removed from the process with streams WW-1, WW-2, WW-3, WW-4 and WW-5. It can be seen from
Table 10 and
Table 11 that these losses are very small. Primarily, the streams are composed of water, the fraction of which ranges from 99.996 mol.% to 99.999 mol.%.
During simulations, it was found that the process is flexible in operation. In cases when surplus power is not available, the process can operate in a hot standby, where small amounts of reactants are consumed. The production of synthetic methane adapts to the hydrogen production. In this case, the process can operate at a lower pressure. When surplus electricity becomes available again, it is used to produce the hydrogen via water electrolysis. The pressure can be increased simultaneously, which expands the capacity of the process. Higher pressure also increases the maximum achievable CO2 conversions, since they are limited by the reaction equilibrium. In cases, where the CO2 content in the final product is too high, its removal would be necessary. The use of higher pressures in our study ensures a sufficiently high methane content. The process, therefore, does not require the upgrading of technologies for CO2 removal, such as water scrubbing, which consumes large quantities of water.
3.4. Energy Streams in the Process Units
Table 12 shows the work and heat flows of the process units for both cases. In the table, we first show the amount of electricity consumed at 100% operating capacity of the three types of electrolysers. The AEC has the highest consumption and the SOEC the lowest, as the latter has the highest efficiency of the three.
With multi-stage compression in the MC-1 and MC-2 units, there is an intermediate cooler between the stages. Its heat flow is expressed as a negative value, because the heat is released to the cooling water. For this reason, in process units where cooling water is present, its consumption is given as a mass flow rate (). As the heat exchangers HE-3, HE-5, HE-6 and CT are also cooled with water, the table shows the negative value of the heat flows and the water consumption for cooling.
For the reactors R-1 and R-2, the negative value represents the heat released in the reaction. The heat is removed from the reactor by the TF flow. In other heat exchangers, the positive values are given for the heat flows. This indicates the heat exchanged between the process streams.
The PTSA is a heat sink. During regeneration, the columns are heated at a water flow rate of 2 kg/s at 100 °C. This flow could be provided with the heat exchanger HE-3, where the water would be heated from 20 °C to 100 °C. With the equation , we can calculate that, in case 1, when less heat is released at the heat exchanger, a water flow of 2.83 kg/s could be guaranteed. In this calculation, we used a temperature change of ΔT = 80 K and a specific heat capacity for water of cp = 4 182 J/(g⋅K). The result of the dynamic simulation of the PTSA unit showed that only 3.99 kW of heat is exchanged during regeneration. The low heat consumption means that the outlet temperature of the heating water is only slightly lower than the inlet temperature. This was confirmed by the average temperature of the heating water at the outlet of the column. On average, the water that passed through the columns was cooled by only 0.49 °C.
There is not much additional potential for heat integration in the process scheme, as the only heat sink in a steady-state operation is the preheating of the reactant. This integration is already ensured with exchangers HE-1 and HE-4. Although PTSA requires heat, the required heat flow is practically negligible compared with the other heat flows in the P2M process.
Instead of heat integration, it is possible to implement cogeneration or simultaneous electricity production. A large part of the heat released in the methanation process is already used to produce superheated medium-pressure steam, which is used to generate electricity in the turbines. In the RC, the CO2 methanation produces 946.67 kW, but when biogas is mixed at the inlet, the work produced is lower, at 938.87 kW. The lower value is to be expected, as the overall CO2 conversion is lower in the case of simultaneous methanation. As a result, overall, less heat is released in the reaction, and the amount of steam produced is therefore lower, but the change does not have a significant impact on the required pump work, which, in both cases, amounts to 1.69 kW.
According to the results in
Table 12, the methanation stage is still a large consumer of cold utilities. Large quantities of cooling water are required for the intermediate cooling in two-stage compressors and in the heat exchangers HE-3, HE-5 and HE-6. To reduce, or even eliminate, the consumption of cold utilities in the methanation section further, this heat could be used to produce hot sanitary water.
There is also the possibility of introducing another level of low-pressure steam in the steam cycle. This would increase the amount of produced electricity. The introduction of an additional level would make use of the heat and shift the cooling water consumption from the methanation section to the steam cycle, where it would be used to condense the saturated steam. It can be seen from the table that the condenser CT, which is used to condense vacuum steam, consumes large amounts of cooling water. The introduction of a new cycle would increase the water consumption further. The released heat cannot be used directly in any other way, as it is low-temperature heat at 31 °C.
Due to the high consumption, the operating costs of the electrolyser can also be very high. To cover even a small amount of the needed electricity, it would be reasonable to improve the co-generation of electricity in further simulations.
As the simultaneous generation of electricity results in an additional investment due to the required process units, it would also be reasonable to carry out an economic analysis of the different process configurations.
It is important to highlight the special case of using SOEC for hydrogen production. The SOEC must be operated at sufficiently high temperatures to ensure a proper electrolysis operation. Heat can be supplied by using the released reaction heat to produce water vapour. The vapour is then used in the electrolysis to produce hydrogen [
8]. The electricity consumption is reduced, because a part of the energy needed for the electrolysis is provided by heat [
19]. For this reason, it would also be worthwhile to investigate the possibility of integrating SOEC and the methanation section of the P2M process further.