The reversible conversion process consists of two separate steps: one for hydrogenation and one for dehydrogenation (as illustrated in Figure 2). In the hydrogenation step, the dehydrogenated carrier is mixed with hydrogen and preheated to 179 °C by cooling the hot stream exiting the hydrogenation reactor. Because of the strong exothermic nature of the reaction, it is proposed that the missing heat to reach the reaction temperature of 200 °C could be supplied by feeding the stream directly into the reactor. The reactor still produces about 50 kJ of waste heat per mol hydrogen, and because of the high temperature, this waste heat could be used for domestic heating, for example. Integration of this heat into the endothermic dehydrogenation process was not considered, because the temperatures do not match, and significant gaps in time can occur between the two process steps. Hydrogenation is carried out at elevated pressures of between 30 and 50 bar. Since electrolysis can be carried out at elevated pressures without a loss of efficiency we did not consider additional energy demand for compression.
In the dehydrogenation step, the hydrogen carrier is preheated (preheater 1 in Figure 2) by cooling the hot liquid stream coming out of the dehydrogenation reactor. A second preheater is used to heat the stream to the reaction temperature of 250 °C. Because of the endothermic nature of this process, additional energy must be supplied to the reactor. Dehydrogenation is carried out at 1 bar. The dehydrogenated carrier is later removed by partial condensation and the hydrogen is converted to water in a fuel cell.
An efficiency of 70 % relative to the lower heating value of hydrogen was assumed for the electrolysis and 55 % for the fuel cell. The heat demand of the dehydrogenation step is 86.0 kJ per mol hydrogen. The overall efficiency of the process thus can be calculated to be approximately 30.8 %. Similar results are obtained for reversible hydrogen conversion using other organic hydrogen carriers. Using toluene as a carrier compound, we calculated the overall efficiency to vary between 30.4 % and 31.2 % (depending on the process used). For azaborines we could show in previous work that an overall efficiency of up to 30.7 % can be reached.4
For the irreversible-conversion storage process, a carbon resource is required, and CO2 instead of CO has been considered here to achieve a CO2-neutral process. An appropriate model must take the differences between these feedstocks into account, particularly the negative effect on catalytic performance, and the influence on the product selectivity.14, 15
The products of Fischer–Tropsch synthesis are very complex. More than a hundred different components, including paraffins, olefins, alcohols, and some aromatic compounds can be contained in the product stream. However, alkanes normally dominate hydrocarbon production in the Fischer–Tropsch process. As a simplification we therefore considered alkanes to be the only products of the reaction, because our focus was not a detailed modeling of the reaction mechanism but of the respective energy balance. Additionally, the energy demand for further byproduct removal is low compared to the total energy demand, and separation is not necessary because traces of alcohols or aromatics in the gasoline are acceptable and can even produce positive effects.
Generally, the product distribution of hydrocarbons formed during the Fischer–Tropsch process follows an Anderson–Schulz–Flory distribution (ASF)16 (see Figure 3), which can be expressed as(1)
Figure 3. Product distribution in Fischer–Tropsch reactions as a function of the chain-growth probability according to Ref. 12.
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with Wn being the weight fraction of hydrocarbon molecules, containing n carbon atoms and α being the chain growth probability.
Although CO2 would be the preferred feedstock for the irreversible chemical storage of hydrogen, the direct conversion from CO2 as the only feedstock to liquid fuels is not the state-of-the-art technology. For this case many experiments show a dominancy of methane production instead of long chain alkane growth.17, 18
To achieve a realizable process, we considered a two-stage process. Carbon dioxide is converted to CO and H2O by using a high-temperature RWGS reactor and the produced gas mixture is fed to the Fischer–Tropsch reactor after removing the water.
In the presence of CO, CO2 is negligibly hydrogenated; it has only a minor effect on product distribution and behaves as an inert species.18 Synthesis using a gas mixture of H2/CO/CO2 has already been researched extensively, because it is also desirable for a conventional—that is, coal based—Fischer–Tropsch process to use the raw feed stream after coal gasification, without employing an additional CO2 separation process.
A description of the process simulation is shown in Figure 4. H2 and CO2 are mixed with recycling streams and compressed to a pressure of 30 bar. The RWGS reaction is carried out in reactor R1, which was modeled by an R-Gibbs Reactor from Aspen Plus. The produced syngas stream is cooled in a heat exchanger block and then flashed in block F1 to remove the water.
A syngas mixture with a CO2/CO/H2 molar ratio of 0.65/2.82/6.52 enters the Fischer–Tropsch Reactor (R2), which is modeled as an R-Yield reactor. Fortran user block models were used, which contained the yield calculation of conversion processes based on the ASF-function [Eqn. (1)].
The chain growth probability (α) is assumed to be 0.9, and the conversion of CO to be 70 %, which is a typical value for the low-temperature Fischer–Tropsch reaction. The product stream contains a hydrocarbon mixture of various alkanes, unreacted CO, H2, and CO2.
The product stream is flashed at different temperatures (F2 and F3). The resulting liquid stream, which consists of a water-rich- and an organic phase, is separated by decantation. The organic phase, with wax as the dominant product, is converted to a heavy-hydrocarbon product downstream in a hydrocrack reactor (R3).
The gaseous product is cleaned by using two separate steps to remove the H2, CO, and CO2. The hydrogen is recovered by pressure swing adsorption (S1). CO and CO2 are separated from the hydrocarbon gas by using a cryogenic treatment at −65 °C.
The hydrocarbons derived from the separation units S2 und D1 are separated in a distillation column (C1) into light- and heavy-hydrocarbon products. A part of the light hydrocarbons (mainly C1–C4) is considered here as exhaust, which could also be condensed by using high pressure or low temperature. However, because they only occur in small amounts, they have not been taken into account. Heat integration has been assumed for the whole process by using a pinch point analysis.
The overall efficiency of the irreversible-conversion process has been evaluated by assuming an efficiency of 70 % for the electrolysis step as before and an energy demand of 2 MJel kg−1 CO2 for the separation of CO2 from the flue gas.7 The efficiency of the irreversible conversion can be calculated as the ratio of the energy output divided by the energy input:(2)
for which the symbols are defined in the section below.
We have modeled the reconversion of the fuel into work using common combustion engines such as those used in cars. The light hydrocarbons (gasoline) are used in an Otto motor with an efficiency of ηgasoline=30 % and the heavy hydrocarbons (diesel) in a Diesel motor with an efficiency of ηdiesel=40 % (both values for full load operation according to Ref. 19). The overall efficiency of the whole process thus can be calculated to be approximately 16.7 %. This is in the same order of magnitude as the efficiency of methanization with reconversion in a gas motor (21.4 %).7 For methane, other options for electricity production are possible (e.g., a combined-cycle power plant) and therefore higher overall efficiencies are conceivable. However, the efficiencies of irreversible conversion processes remain lower than those of reversible conversion.