Design Study Report
ANTECY solar fuels development
Project runtime:
December 2014 – August 2015
Team members:
Alexander van der Made (Shell)
Robert Moene (Shell)
Tim Nisbet (Shell)
Saša Marinić (ANTECY)
Paul O’Connor (ANTECY)
Timo Roestenberg (ANTECY)
Subject:
Technical and economic evaluation of the ANTECY solar fuels process
Report author:
Timo Roestenberg (ANTECY)
Report date: 07-09-2015
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Table of contents 1 Executive summary ......................................................................................................................... 3
2 Project summary ............................................................................................................................. 5
3 Introduction ..................................................................................................................................... 7
3.1 Goal of the design study .......................................................................................................... 8
3.2 Approach ................................................................................................................................. 8
4 Background ...................................................................................................................................... 8
4.1 The technology ........................................................................................................................ 8
5 Project findings .............................................................................................................................. 11
5.1 Technical evaluation .............................................................................................................. 11
5.2 Economic evaluation ............................................................................................................. 12
6 Future perspective ........................................................................................................................ 13
Appendix I: detailed cost estimate of the 5 MW case .......................................................................... 17
6.1 Estimation of CAPEX of CO2 sequestration step .................................................................... 18
6.2 Estimation of CAPEX of hydrogen production by electrolysis ............................................... 22
6.3 Estimation of CAPEX of methanol synthesis ......................................................................... 24
6.4 Total main equipment cost ................................................................................................... 25
6.5 Real construction cost ........................................................................................................... 26
6.6 Discounting cost .................................................................................................................... 26
Appendix II: detailed cost estimate of the 500 MW case ..................................................................... 29
Appendix III: projected cost of PV and green hydrogen ....................................................................... 32
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1 Executive summary
The technical and economic feasibility of converting renewable energy produced by Solar (and or
Wind) into liquid fuels and/or chemicals were evaluated in a joint design study for a semi-commercial
and commercial plant. This study was performed by staff from Shell1 and ANTECY2.
The technology required comprises the following steps:
A. Production of electricity from solar energy via photovoltaics.
B. Production of hydrogen by water splitting (Electrolysis).
C. Capturing and concentrating CO2 from Air (and/or Flue gas).
D. Conversion of CO2 and hydrogen into methanol.
Steps A, B and D are based on state-of-the-art proven technology, while the technology for Step C is
under development by ANTECY and has been demonstrated (“Proof-of-Concept”) at a laboratory
scale.
The CO2 capturing and concentration process developed by ANTECY has certain distinct benefits over
competing technologies:
Reduced energy consumption by improved heat integration between the steps enabled by
the low desorption temperature of the adsorbent system, allowing the use of low value heat.
The use of an environmentally friendly non-degrading solid adsorbent system.
With regards to the technical feasibility of the overall process and the specific new technology to
capture and concentrate CO2, no technical barriers are foreseen to starting process development for
implementation at a semi-commercial (5 MW, approx. 2,6 kton methanol/year) and commercial
(500 MW, approx. 260 kton methanol/year) scale.
With regards to the economic feasibility at the present state of the technology and assuming a low
(but not zero) marginal electricity value of 0.02 € /kWh a 500 MW commercial plant would not yet be
profitable.
But the expectation is that the following developments can lead to improved economics within a 10
to 25 year timeframe, being:
1. An expected strong reduction of the electrolyzer CAPEX costs.
2. An increase of the overall energy efficiency of the process (Steps A to D) from ± 50% to 60%.
The implementation a CO2 tax, of ± 40 €/ton CO2 or more can also have a significant positive effect
on the overall economics.
The above is not yet taking into account further anticipated improvements in electrolyzer energy
efficiency and the closer integration of electrolysis with solar energy captured either by super-
photovoltaics (increasing solar energy capturing efficiency from ±20% to 40%) or other improved
forms of solar energy concentration.
1 Shell Global Solutions International B.V., Amsterdam
2 Antecy BV, Hoevelaken.
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The interplay between scale-up and cost reductions in electrolysis and CO2 capturing is expected to
lead to a point where CO2 capturing and utilization with renewable energy (electrons) becomes
interesting.
Therefore there is a sufficient incentive to further develop and scale-up this technology and
demonstrate the critical parts on a pilot and demo scale. The recommendation is for a next step to
set up a project team to address this. ANTECY will lead this effort while the proposal is that Shell will
continue to participate and support the project in the project definition phase, by jointly establishing
the risk factors, and the scale and sequence of demonstration and de-risking.
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2 Project summary
This report outlines the findings of the six month design study that was executed in a joint effort
between ANTECY B.V. and staff from Shell Global Solutions International B.V. The goal of the study
was to evaluate the technical and economic feasibility of the Solar Fuels process under development
by ANTECY, and to pinpoint those areas that merit the most improvement efforts. The Solar Fuels
process under development by ANTECY sequestrates CO2 from ambient air, splits water in hydrogen
and oxygen by means of electrolysis, and reacts the hydrogen and CO2 in a thermocatalytic reactor to
methanol, using renewable energy. Renewable energy is thus stored in a chemical that can be
integrated in the existing fuel and carbon-chemicals infrastructure.
The approach for the design study was to consider two cases. Firstly, a design and economic
evaluation was done for a process that has a power input of 5 MW. This case was evaluated based on
technical feasibility and scalability. The economic evaluation was done by obtaining estimates for the
main equipment, which was then multiplied by realistic multiplication factors that are common
practice at Shell, to arrive at a total capital investment cost. Based on the 5 MW case (approx.
2,6 kton methanol/year), a translation step was made to a 500 MW (approx. 260 kton
methanol/year) case using customary scale up rules. The results of the translation were cross
checked to figures of known installations of comparable processes at similar scale.
The main conclusions of the study are as follows.
With respect to the technical feasibility: key innovations made by ANTECY in this process are the CO2
sequestration and the energy integration between steps of the process. Each of the process steps
were individually evaluated on technical feasibility, as well as their integratability.
Renewable energy generation at both small and large scale has been demonstrated. Both
efficiencies and economics are improving, and expected to keep doing so over the years to
come. The footprint of the Solar Fuels process itself has been calculated to be 10-20 % of the
associated required large scale solar power generation system.
The CO2 sequestration process developed by ANTECY is deemed technically feasible, with no
technical barriers to development identified. It is also the process step that will require the
most effort to bring from current development state to industrial application. It is also the
only step that in the process that is entirely new, all other process steps have been proven on
industrial scale and are readily available.
Water electrolysis is a mature technology that can be bought off the shelf. Still progress is
needed (and expected) in terms of energy efficiency and, in particular, CAPEX, in order to
improve the economics of Solar Fuels production.
Methanol synthesis from CO2 and H2 is also a mature technology possible with commercially
available catalysts and reactor designs. Selectivity and activity of the process have been
shown to be competitive with alternative routes.
The mass and energy flows of the different process steps have been calculated to make
integration into a viable process possible. Heat generated by the hydrogen production and
methanol synthesis can supply the heat required for the CO2 sequestration step.
Methanol is deemed to be a good target product for a Solar Fuels process, because it forms a
dense and easily integratable energy carrier in the current energy infrastructure and because
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it forms a platform chemical for the production of carbon-chemicals such as olefins and
aromatics.
In conclusion it was found that no technical barriers are foreseen by the team to starting process
development for the implementation of the technology on a 5 MW to 500 MW scale.
With respect to the economic feasibility: a cost estimation was performed for the Solar Fuels process
at 5 MW and 500 MW scale. In the estimation current costs for equipment were used and a flat rate
of 0.02 €/kWh3 was used for energy required, with an assumed availability of 24h per day. For the
process to be economically profitable the cost of hydrogen production by electrolysis needs to be
reduced. This is illustrated in the charts in Figure 1 and Figure 2 where it can be seen that, especially
in the large scale case, the electrolyser cost is a substantial part of the total cost. It is projected that
the required improvements in electrolysis economics may occur in the 2025 and beyond timeframe.
It is the conclusion of the team that the results of the design study justify piloting the technology at
this time. The pilot should focus on the novel aspects of the CO2 capture technology to test the
validity of the assumptions made in this design study. The recommendation is for a next step to set
up a project team to address this. ANTECY will lead this effort while the proposal is that Shell will
continue to participate and support the project in the project definition phase, by jointly establishing
the risk factors, and the scale and sequence of demonstration and de-risking.
Figure 1, Cost breakdown of the 5 MW case (blue: hydrogen production, red shades: methanol synthesis, green shades: CO2 capture), total CAPEX 14.2 mln€.
3 Current (2014) bulk, grid energy price in the US is approx. 0.04 USD/kWh.
29%
11%
4%
25%
2%
4%
12%
6% 7%
CAPEX 5 MW
Electrolyzer
H2 compressor
CO2 compressor
Reactor
Vacuum pump
Desorber tank
Blower
Adsorber
Sorbent
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Figure 2, Cost breakdown of the 500 MW case (blue: hydrogen production, red shades: methanol synthesis, green shades: CO2 capture), total CAPEX 748 mln€.
3 Introduction
This report outlines the findings of the six month design study that was executed in a joint effort
between ANTECY B.V. and staff from Shell Global Solutions International B.V. The goal of the study
was to evaluate the technical and economic feasibility of the Solar Fuels process under development
by ANTECY, and to pinpoint those areas that merit the most improvement efforts.
The Solar Fuels process under development by ANTECY sequestrates CO2 from ambient air, splits
water in hydrogen and oxygen by means of electrolysis, and reacts the hydrogen and CO2 in a
thermocatalytic reactor to methanol, using renewable energy. Renewable energy is thus stored in a
chemical that can be integrated in the existing fuel and energy infrastructure, as well as form the
basis for a range of green chemicals.
For the study, two main design cases have been considered for the solar fuels process. Firstly, a
design and economic evaluation was done for a process that has a power input of 5 MW (approx.
2,6 kton methanol/year4). This case was evaluated based on technical feasibility and scalability. The
economic evaluation was done by obtaining estimates for the main equipment, which was then
multiplied by realistic scaling factors that are common practice at Shell, to arrive at a total capital
investment cost.
Based on the 5 MW case, a translation step was made to a 500 MW case (approx.
260 kton methanol/year) using customary scale up rules. The results of the translation were cross
checked with figures of known installations of comparable processes at similar scale.
Finally, a technical and economic evaluation was done for the CO2 capture part of the process,
applied to flue gasses, specifically, compared to state of the art amine processes.
4 The energy equivalent of one ton of methanol is approx. 3.2 barrels crude oil.
48%
5%
2%
10% 1%
7%
5%
10%
12%
CAPEX 500 MW
Electrolyzer
H2 compressor
CO2 compressor
Reactor
Vacuum pump
Desorber tank
Blower
Adsorber
Sorbent
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3.1 Goal of the design study The goal of the design study is to evaluate the technical and economic feasibility of the ANTECY solar
fuels concept, today and in the future. For the economic feasibility an accuracy target of ±50% was
desired.
3.2 Approach To achieve this evaluation, the technical design was discussed and, where needed, modified. For
each of the main components, cost estimates were obtained. These were used for a CAPEX estimate.
This was done for a 5 MW size installation. Subsequently, scale up rules of that are common
engineering practice were applied to estimate the CAPEX of a 500 MW size installation. The results of
the translation were cross checked with figures of known installations of comparable processes at
similar scale.
The reason a 500 MW size for the calculation was chosen is that this was deemed to be a realistic size
to be compatible with large scale sustainable energy generation. Typical large scale sustainable
energy generation systems, such as large scale solar PV and large scale concentrated solar power
plants, that have been constructed and are planned for the coming years, are in the 500 MW scale
range. However, the results obtained for the 500 MW case can be extrapolated to larger sizes using
the same scale up rules outlined in this report.
4 Background
4.1 The technology The ANTECY solar fuels technology consists, in its simplest form, of three main steps. Firstly, CO2 is
sequestrated from either air or a different CO2 containing gas stream (such as flue gas). At the same
time, water is split electrochemically to form hydrogen and oxygen gas. The hydrogen gas is reacted
with the CO2 in a thermocatalytic reactor to form methanol5. In this way, sustainably generated
electricity is stored in the form of a liquid fuel. This is schematically shown in Figure 3.
Additionally, the oxygen produced could be monetized to increase the profitability of the process,
but this was not included in the scope of this design study.
Of the technology, the latter two steps can be considered “state of the art” and technically proven;
both water electrolysis and CO2 hydrogenation to methanol are applied on varying scales. The CO2
sequestration step is the most novel, and at the same time the step that ANTECY has spent most of
its efforts on developing.
The CO2 sequestration process developed by ANTECY, named CAIR, comprises a sorbent (a hydrated
form of potassium carbonate on a support) and a specific process developed to efficiently use the
sorbent. In Figure 4 the process for producing pure CO2 from a CO2 containing gas is schematically
shown. The process is a cyclic process, where the CO2 containing gas is contacted with the sorbent
under ambient conditions. The sorbent is then heated under reduced pressure in the presence of a
flow of water vapour. The water vapour serves as a scavenging gas to retrieve the released CO2. The
CO2 can be easily separated from the water vapour by condensation.
5 Methanol is just one example of a potential product; alternative products include Fischer-Tropsch liquids.
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Figure 3, Schematic representation of the solar fuels process.
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Figure 4, Schematic illustrating the CAIR process principle.
There are variations possible to the CAIR principle as illustrated. For example, it may be
advantageous to have the adsorption and desorption taking place in different reactors. In this way,
the adsorption reactor can be optimized for contacting the CO2 containing gas with the sorbent,
under ambient conditions, while the desorption reactor is designed for the desorption process, at
higher temperature and reduced pressure. In this way one desorber can be used to regenerate the
sorbent of multiple adsorption reactors. This can, in the case of for example CO2 adsorption from air,
reduce capital cost. An additional advantage is that the height/width ratio of the adsorption reactor
can be made very low (in the order of 1/40), reducing pressure drop and thus reducing operational
cost. This example is illustrated in Figure 5. Further details on the dimensions are available in:
Appendix I: detailed cost estimate of the 5 MW case.
Figure 5, Schematic of the CAIR process using two reactors.
The sequestration of CO2 is done by means of a solid salt that is deposited on a carrier. Specifically, a
hydrated form of potassium carbonate is used. The chemical reaction that takes place is:
𝐶𝑂2(𝑔) + 𝐾2𝐶𝑂3 ∙ 1.5𝐻2𝑂(𝑠) ↔ 2𝐾𝐻𝐶𝑂3(𝑠) + 0.5𝐻2𝑂(𝑔)
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In this reaction the hydrated form of potassium is potassium carbonate sesquihydrate. However,
there are many other hydrated forms of potassium carbonate possible, depending on the conditions
of the reaction.
The advantage of the reaction with a hydrated form of potassium carbonate, as opposed to an
anhydrous form, is twofold. Firstly, the reaction enthalpy is lower, meaning that the amount of
energy that needs to be spent to reverse the reaction (desorb CO2 from the sorbent) is lower, thus
also reducing the reaction reversal temperature6. Secondly, since water is a product of the
adsorption reaction, and a reactant for desorption (as opposed to the opposite when using the
anhydrous reaction mechanism), the equilibrium of the desorption of CO2 can be shifted by
desorbing in a moist environment.
The footprint of the process, as described above, is a factor of 10-20 % of the footprint of the large
scale solar power generation plant of the same capacity7. Since it is likely that not the complete
capacity of the solar power generation will be used for solar fuels production, this will make the
required footprint of the solar fuels system relatively even smaller.
5 Project findings
5.1 Technical evaluation The solar fuels production process consists of four main components: energy generation, CO2
capture, hydrogen generation and methanol synthesis. The entire process has been simulated in a
flow sheet, calculating mass and energy flows of all mayor components. The technical evaluation of
the four main process steps will be discussed separately.
Energy generation
The energy required for the process can be generated by various means, all of which are already
applied on industrial as well as some on residential scale. Photovoltaics have been improved greatly
over the last decades, both in efficiency as well in production costs. The same can be said for wind
turbines. Large scale solar thermal plants are also already applied, although only a handful exist
around the world currently. So, more progress in large scale CSP (Concentrated Solar Power) can be
expected in the future. Since there is a multitude of options for energy generation that can be
combined with the Solar Fuels concept, each with their own merits, drawbacks and costs, a flat rate
of 0.02€/kWh was assumed for the cost of electricity in the calculations of this design study.8
Appendix III: projected cost of PV and green hydrogen, presents the base case and more aggressive
scenarios we used for costs of solar PV electricity up to 2050.
The various options with their varying levels of maturity do not inhibit the feasibility of the
application of Solar Fuels. In fact, the increasing market penetration of sustainable energy sources
and associated cost reductions only adds to the viability of Solar Fuels.
6 For the reaction shown, the reaction enthalpy is -41 kJ/mol, while the reaction enthalpy of the corresponding
anhydrous potassium carbonate reaction is -141 kJ/mol. 7 The land use of solar PV power generation is about 10,000 m
2/MW (1 GW plant, 20% solar cell efficiency),
while the CO2 capture, which constitutes the bulk of the land use of the Solar Fuels process, is approx. 1100 m
2/MW.
8 NB: the flat rate of 0.02 €/kWh includes the cost of storage of intermittent solar PV power, if PV is used.
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Energy intermittency
One key aspect of many sustainable energy generation options is their intermittent nature. This
intermittency manifests itself from rapid (order of minutes-hours) fluctuations, to slow (weeks-
months) fluctuations. While the Solar Fuels is ideally suited to be used as a long term energy storage
(and can thus be employed to solve problems with slow fluctuations), short fluctuations effect the
economics of the process. In order to meet the short term fluctuation, the Solar Fuels technology can
be combined with a short term energy storage solution. Since there are many options for short term
energy storage, such as batteries, compressed air energy storage, capacitors, molten salts. The ideal
solution is dependent on the application location, the choice for short term energy storage type is
left out of the scope of this design study, and assumed to be included in the future cost of electricity.
CO2 capture
The CO2 sequestration step is the most novel, and at the same time the step that ANTECY has spent
most of its efforts on developing. For this reason, the CO2 sequestration technology has been
analysed in most detail within this design study. During the different meetings a novel concept was
put forth for the application of the technology as developed by ANTECY. Based on this new concept it
was concluded that no technical barriers are foreseen for the implementation of the CO2
sequestration technology. Some questions remain regarding kinetics and full process integration, so
a demo is advocated.
Hydrogen generation
The concept of water splitting into hydrogen and oxygen by electrolysis has been known for over two
hundred years since the first experiments by William Nicholson and Anthony Carlisle around 1800.
While the technology is close to technical maturity, significant improvements in the economics of the
large scale application are still expected. For the technical feasibility no problems are foreseen, for
the economic evaluation of the solar fuels process, the current state of the art is used, with the side
note that great improvement is expected over the coming years. Appendix III: projected cost of PV
and green hydrogen presents the base case and more aggressive scenarios we used for costs of
electrolyser CAPEX and corresponding costs of hydrogen up to 2050.
Methanol synthesis
Methanol synthesis by CO2 hydrogenation, or, alternatively, by CO2 reduction to CO by RWGS
(Reverse Water Gas Shift), followed by CO hydrogenation, are both applied by industry already9.
Commercial catalysts and reactor designs are available for both routes. Selectivity and activity are
comparable for both routes. Tests with commercially available catalysts have been carried out by
ANTECY, achieving selectivity of 99,5%, at an estimated conversion of 25-30% per pass, and activity
of 6 mmol MeOH/g cat/hr. For the solar fuels process, the direct hydrogenation of CO2 is deemed
most feasible. The technology is currently mature enough to be applied in the solar fuels process
without foreseen technical barriers.
5.2 Economic evaluation For each of the main components, cost estimates were obtained. These were used for a CAPEX
estimate. This was done for a 5 MW size installation. Subsequently, scale up rules were applied to
9 Direct CO2 hydrogenation is, for example, done by Mitsui chemicals in a 100 ton/year pilot, constructed in
2008. The more classical route via syngas through RWGS is, for example, done by Audi with plant operator Sunfire in Dresden.
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estimate the CAPEX of a 500 MW size installation. The details of the economic evaluation can be
found in Appendix I: detailed cost estimate of the 5 MW case and Appendix II: detailed cost estimate
of the 500 MW case. The economic evaluation was done for a 25 year lifetime of the plant, with
different discount rates. The results of the economic evaluation, as projected at the current level of
technological progress, and present day prices, for both the 5 MW and 500 MW case are shown in
the following table. In this table, both the CAPEX, the total cost discounted over 25 years and the
methanol production cost (calculated as total discounted cost / discounted tons of methanol
produced) are shown. Details of the cost estimates can be found in: Appendix I: detailed cost
estimate of the 5 MW case, and: Appendix II: detailed cost estimate of the 500 MW case.
5 MW 500 MW
CAPEX (2015) 14.2 M€ 748 M€
Total discounted cost (10% discount rate) 20.9 M€ 1349 M€
(15% discount rate) 18.6 M€ 1154 M€
Discounted cost / discounted ton methanol
(10% discount rate) 861 €/ton 555 €/ton
(15% discount rate) 1089 €/ton 675 €/ton
6 Future perspective
The two key points for the future perspective of the technology that have been identified are the
future energy cost and the cost of hydrogen production by electrolysis. The cost of electricity
production by PV has been declining steadily for decades. As PV becomes more popular, market
share becomes higher and technological progress continues, this trend is expected to continue, as is
shown in Appendix III: projected cost of PV and green hydrogen. However, the solar fuels technology
can be applied with any (mix of-) electrical energy available. At the moment, depending on location,
bulk industrial energy price is around or below 0,04 €/kWh10. In the Netherlands, current bulk
industrial energy price is 0,017 €/kWh (excluding taxes and tariffs)11. Thus, the assumption of an
energy cost of 0,02 €/kWh as a future bulk energy price seems not unrealistic.
The market for hydrogen production by electrolysis is still not well developed. The learning curve for
the large scale deployment of the technology is yet to be fully mounted. As hydrogen production by
electrolysis is deployed more and more, costs are expected to decline significantly. In optimistic
scenarios the cost could be cut to a third in the next decade, in less optimistic scenarios this would
take 25 years (see Appendix III: projected cost of PV and green hydrogen). Since the hydrogen
production by electrolysis currently constitutes almost 50% of the total investment for the 500 MW
case (360 mln€ of the 748 mln€ total), this is expected to have a very pronounced positive effect on
the economics of the technology over the coming years.
Furthermore, ongoing development of the process by ANTECY will further enhance the viability of
the process. At the current stage of development, the electric energy use of the CO2 separation
process is approx. 5,5 times the theoretical energy required (±2500 kJ/kg needed in the Solar Fuels
process vs 450-500 kJ/kg, depending on conditions, in theory). The total energy use (heat and
10
State energy data system, Energy Information Agency, http://www.eia.gov/state/seds/seds-data-complete.cfm, accessed last on 07-09-2015. 11
Kathleen Jennrich, Katharina Grave, Dr. Barbara Breitschopf, European electricity prices and their components, published by Ecofys Germany, Fraunhofer Institute, and the German Ministry of Economic Affairs and Energy, July 2015.
antecy 14
electricity combined) is about 22 times the theoretical energy requirement for separation. Thus,
there is room for improvement, and an increase of total system efficiency from 50%12 to 60% is
deemed realistic.
To illustrate the effect of the advances in both the technology as well as its components, a future
perspective has been created for the following cases:
Electrolyser cost is reduced by 50%
Process efficiency is increased to 60%
Both electrolyser cost halves and process efficiency is increased to 60%
The result for the 500 MW case is shown in the following table (“discount rate” has been abbreviated
to “d.r.”).
Current Electrolyser cost 50% Process η 0.6 Both
CAPEX (future) 748 M€ 568 M€ 748 M€ 568 M€
Total discounted cost (10% d.r.) 1349 M€ 1175 M€ 1349 M€ 1175 M€
(15% d.r.) 1154 M€ 982 M€ 1154 M€ 982 M€
Discounted cost / discounted ton methanol
(10% d.r.) 555 €/ton 483 €/ton 461 €/ton 402 €/ton
(15% d.r.) 675 €/ton 573 €/ton 561 €/ton 477 €/ton
Another aspect that can be considered for the future application is the possible value of CO2 taxation.
The effect of a CO2 tax of 40 and 80€/ton on the process economics is presented below:
Current CO2 tax at 40 €/ton CO2 tax at 80 €/ton
CAPEX 748 M€ 748 M€ 748 M€
Total discounted cost (10% d.r.) 1349 M€ 1349 M€ 1349 M€
(15% d.r.) 1154 M€ 1154 M€ 1154 M€
Discounted cost / discounted ton methanol
(10% d.r.) 555 €/ton 500 €/ton 445 €/ton
(15% d.r.) 675 €/ton 620 €/ton 565 €/ton
In Figure 6 the different possible scenarios (advances in technology and CO2 taxation) are calculated
for different combinations of possible cases. The following scenarios are depicted (in order of
decreasing net production cost):
2015: 1 - The situation as is
2020: 2 - CO2 taxation at 40 €/ton
3 - Process η to 60%
2030: 4 - Electrolyser cost halves and CO2 taxation at 40 €/ton
5 - Electrolyser cost halves and CO2 taxation at 80 €/ton
12
Process efficiency is defined here as the ratio between chemical energy in the product and total electric energy into the process. The current design is evaluated to have an efficiency of approximately 50%.
antecy 15
6 - Process η to 60% and CO2 taxation at 40 €/ton
7 - Process η to 60% and CO2 taxation at 80 €/ton
8 - Electrolyser cost halves and CO2 and Process η to 60%
2040: 9 - Electrolyser cost halves and CO2 taxation at 40 €/ton
10 - Electrolyser cost halves and CO2 and Process η to 60% and CO2 taxation at 40 €/ton
11 - Electrolyser cost halves and CO2 and Process η to 60% and CO2 taxation at 80 €/ton
From the graph it can be seen that the dependence on discount rate is heavy, which is caused by the
process being very capital intense. However, most scenarios predict that, at a discount rate of 10%, it
will be possible to produce methanol for a competitive price. An added effect, which was not taken
into account in this calculation is the price evolution of commercial methanol, beyond inflation,
which is likely to have an additional positive effect on the competitiveness of the Solar Fuels process.
This is shown in Figure 7. The projected methanol price is based on the calculation shown with Figure
10 (5% yearly), but removing the effect of inflation (a yearly price increase of 2.5% is used).
Figure 6, Current and future cost of methanol production.
1
3
2
6 5
8
4
7 10
11
9
1
3
2
6 5
8
4
7
10
11
9
200
250
300
350
400
450
500
550
600
650
700
2015 2020 2025 2030 2035 2040
Dis
cou
nte
d c
ost
/dis
cou
nte
d t
on
(€
/to
n)
Year of implementation
Future discounted cost/discounted ton
10% discount rate
15% discount rate
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Figure 7, Current and future cost of methanol production compared to projected methanol price.
In Figure 8 the future cost of methanol calculated in barrel oil equivalent (on energy basis) is
displayed.
Figure 8, Current and future product cost per Barrel of Oil Equivalent.
200
250
300
350
400
450
500
550
600
650
700
2015 2020 2025 2030 2035 2040
Dis
cou
nte
d c
ost
/dis
cou
nte
d t
on
(€
/to
n)
Year of implementation
Future discounted cost/discounted ton
Conservative 10%
Optimistic 10%
Conservative 15%
Optimistic 15%
Projected methanol price
50.00
70.00
90.00
110.00
130.00
150.00
170.00
190.00
210.00
2015 2020 2025 2030 2035 2040
Co
st/B
OE
(€/B
OE)
Year of implementation
Cost/BOE
Conservative 10%
Optimistic 10%
Conservative 15%
Optimistic 15%
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Appendix I: detailed cost estimate of the 5 MW case
CO2 capture
Of the process the CO2 production has undergone the most developed efforts by ANTECY and the
concept that is calculated in this document is as follows:
During the meeting of 29-01-2015 a concept was discussed for the adsorption of CO2 from air,
whereby the solid adsorbing particles move through a bed that is flushed in counter flow with air.
Upon exiting the bed, the particles are transported to a vacuum desorption reactor, where the CO2 is
stripped and thereby the particles regenerated. The particles are then returned to the bed, where
CO2 adsorption recommences. This approach has economic benefits, as it minimizes the number of
vacuum vessels needed. During the meeting of 20-03-2015, a modification of this concept was
discussed, whereby the bed does not move during adsorption, and an extra bed is installed. The beds
are flushed with air until they are saturated. The saturated sorbent is then moved into a desorber,
and filled with regenerated adsorbent from the same desorber. By staggering the loading and
regenerating of the different adsorbers, one bed is filled/emptied while the others are being flushed
with air. In this way, the desorber is in continuous use, desorbing the sorbent from one adsorber at a
time.
The global operational parameters of the system were calculated for different parameter choices,
based on this calculation the following case was chosen:
Adsorber:
Bed height: 0.5 m
Adsorber diameter: 20 m
Particle size: 5 mm
Number of adsorbers: 5
Desorber:
Height: 10 m
Diameter: 3.8 m
Assumptions
The solid particles used for the CO2 adsorption are assumed to:
Be free flowing
Be attrition resistant
Have a CO2 adsorption capacity of 10 Nm3 CO2 per m3 of sorbent bed
antecy 18
Concept schematic
Figure 9: Concept schematic (side view of one adsorber and one desorber)
6.1 Estimation of CAPEX of CO2 sequestration step The CO2 adsorbers are a part that will be custom made for this process. As far as possible, off the
shelf parts will be used for the units. To estimate the cost of the vessels an estimate by mass is made
by calculating size of various parts and components. To estimate cost of desorbers, the minimum
wall thickness is calculated
Desorber wall thickness
Calculation of vacuum vessel wall thickness according to ASME VIII Div I Sec A UG-2813: Thickness of
shells and tubes under external pressure (units are inches, psi etc.).
Symbols used:
A = factor determined from chart G in ASME SECTION II PART D14 and used to enter the applicable
material chart in ASME SECTION II PART D.
B = factor obtained from applicable material chart from ASME SECTION II PART D
Do = outside diameter of vessel
E = elasticity modulus at (maximum) design temperature
L = total length of the vessel
P = external pressure (for full vacuum P = 15 psi)
Pa = calculated maximum allowable external pressure
t = vessel wall thickness
13
Part UG, General requirements for all methods of construction and all materials (REFC01) 14
2007 ASME boiler and pressure vessel code, II part D, Properties (Customary) Materials (REFC02)
antecy 19
For the CO2 adsorption vessels the following values are used:
Do = 523 cm
t = 17.3 mm
L = 759 cm
This makes L/D0 = 1.45 and D0/t = 303. Using these values it is found from chart G in ASME SECTION II
PART D, that A = 0.000175. Using the chart for carbon or low alloy steel (fig CS-1 in ASME SECTION II
PART D) the value the maximum allowable external pressure is calculated using:
𝑃𝑎 =2𝐴𝐸
3(𝐷𝑜𝑡
)
For the obtained values of A, D0 and t, and using E = 29e6 (up to 300 F), this gives a Pa of 11.2 psi.
Price per desorber
At a wall thickness of 20 mm, the total volume of steel required for the construction of a vessel of the
dimensions given equals:
Cylinder:
𝑉𝑐 = 𝐿 ∙ 𝑝𝑖 ∙ (𝐷𝑜
2
4−
(𝐷𝑜 − 2𝑡)2
4) = 7.59 ∙ 3.141 ∙ (
5.232
4−
(5.23 − 0.0134)2
4) = 2.15 𝑚3
Top & bottom (together one complete sphere):
𝑉𝑠 =4
3𝑝𝑖 ∙ (
𝐷𝑜3
8−
(𝐷𝑜 − 2𝑡)3
8) = 1.48 𝑚3
𝑉𝑡 = 𝑉𝑐 + 𝑉𝑠 = 3.63 𝑚3
The mass per adsorber comes to (using 7800 kg/m3), 28673 kg.
antecy 20
At a steel price of 1.5 €/kg15, the cost per vessel comes to 1.5*28673 = 43000 €.
Cost of adsorbing material
Total volume needed for 5 MW system is 5 times 0,5 m bed, 20 m diameter, equals 785 m3 sorbent.
Density assumed to be approx. 400 kg/m3, require 314 tons. Assuming price of 2 €/kg, cost is 628 k€.
Blowers
To calculate the pressure drop in the adsorber bed, the Ergun equation is used.
With:
µ = fluid viscocity (dynamic), used is 1.85·10-5 kg/ms.
ε = Void fraction. For a poured random packing of spheres this is 0.375-0.391, for a close random
packing this is 0.359-0.375. Used is 0.375.
Dp = particle diameter, varied between 0.5 and 15 mm.
L = bed height (HA in Figure 9), varied from 0.1 to 3 m.
ρ = fluid density, used is 1.29 kg/m3.
vs = superficial velocity, follows from chosen residence time and bed height.
The chosen parameter values for number of adsorbers and adsorber size determine the residence time. This in turn, together with the chosen particle size determines the pressure drop across the adsorber. Using: µ = 1.85·10-5 kg/ms.
ε = 0.375 -
Dp = 5 mm.
L = 0.5 m
ρ = 1.29 kg/m3.
vs = superficial velocity, follows from chosen residence time and bed height and required airflow, for
the 5 MW case this is 0.32 m/s.
This results in a pressure drop of 400 Pa.
Canada blower has a price list available for centrifugal fans of various sizes (REFC05) (though admittedly not the most energy efficient alternative). The total flow required is 402 m3/s or roughly 850000 SCFM. The largest blowers available delivers a flow of 80000 SCFM, requiring 11 total. The
15
Breeveld staal prijslijst 29-01-2014 (REFC03)
antecy 21
price quoted on the site of Canada blower for this type of fan is roughly 28950 USD (depending on the required head). This brings the cost estimation for the blowers to 318 kUSD or 236 k€. To account for the fact that more energy efficient blowers will be more expensive, but also more economical on the whole, a cost estimation of 300 k€ is used.
Vacuum pump
The concept is to flush the sorbent with dry steam to displace any air in the adsorber. At 20 oC, this
requires a pressure of 25 mbar. In order to displace the air in the reactor with steam, it is assumed
that a volume of dry steam of twice the volume of the reactor is required (more is unlikely to be
prudent, since the absence of CO2 during evacuation will start to cause CO2 to desorb). The flow
required to achieve this is 717 m3/hr.
Additionally, a flow of steam is needed to transport the produced CO2 from the reactor (lowering the
CO2 concentration, allowing more CO2 to be released). The pressure of this flow can be much higher,
since the adsorber will be heated to release the CO2. In order to calculate the flow of steam needed
for this, data from16 is used (specifically, the temperature dependent Gibbs free energy change of
reaction is obtained).
The reaction taking place is:
𝐶𝑂2(𝑔) + 𝐾2𝐶𝑂3 ∙ 1.5𝐻2𝑂(𝑠) ↔ 2𝐾𝐻𝐶𝑂3(𝑠) + 0.5𝐻2𝑂(𝑔)
For any general reaction 𝛼𝐴 + 𝛽𝐵 → 𝜎𝑆 + 𝜏𝑇 it can be said that 𝐾𝑐 =[𝑆]𝜎[𝑇]𝜏
[𝐴]𝛼[𝐵]𝛽. So, in this case:
𝐾𝑐 =√[𝐻2𝑂]
[𝐶𝑂2]
Using ∆𝑟𝐺𝑜 = −𝑅𝑇𝑙𝑛(𝐾) and rewriting, the equilibrium concentration of CO2 can be calculated for a
given concentration of water vapour. The Gibbs free energy of reaction is taken from figure 6 of16 to
be:
∆𝐺 = 88.033 ∙ 𝑇 − 39230
Assuming the reactor is heated to 80oC. The saturation pressure of water at this temperature is
calculated (as before) using the Antoine equation:
𝑙𝑜𝑔10(𝑃) = 𝐴 −𝐵
𝐶 + 𝑇
With A=10.196; B=1730.63 and C=-39,724. (P in Pa and T in K).
For T=353 K, this gives 47 kPa or 0.47 bar. This is thus the highest concentration of water vapour
possible at this temperature without condensation. However, to achieve such a concentration, a
significant amount of heat is needed to generate this water. Thus, a tradeoff can be made between
desorption pressure (energy use of and size of vacuum pump) and heat use. At a vapour pressure of
0.2 bar and a temperature of 80 oC, the equilibrium concentration of CO2 is 0.028 bar. Thus, at this
pressure a flow with a composition 12.3% CO2 and 87.7% H2O and a total pressure of 0.228 bar is
16
Ab Initio Thermodynamic Study of the CO2 Capture Properties of Potassium Carbonate Sesquihydrate, K2CO3·1.5H2O; Yuhua Duan et al; The Journal of Physical Chemistry C 2012 116 (27), 14461-14470 (REFC07)
antecy 22
produced. The heat required to evaporate this flow of water is less than the heat generated by the
hydrogen electrolyser (850 kW of heat needed, more than 1 MW of heat generated).
The total flow that needs to pumped by the vacuum pump in this case is approx 12600 m3/hr, but at
a pressure of 0.228 bar.
A quote was obtained from Pfeiffer vacuum for a pumping station that can pump 35000 m3/hr at a
pressure of 25 mbar. This unit is (at least) three times as big as required, but additionally the data
provided assumes it operates at 25 bar. To get an idea of the change of energy use at higher
operating pressure, REFC0817 is used (datasheet to a different vacuum pump). In the graphs in this
datasheet it can be seen that in the range of 25mbar to 0.228 bar, the flow rate at constant power is
either constant or increasing with increasing pressure.
The price of the Pfeiffer system is 695 k€, and this system is three times too big, and capable of
operating continuously at a too deep vacuum. For this reason an estimate of 50 k€ is used.
Adsorber
To estimate the cost of the adsorbers, a breakdown into likely parts was made. These are:
- Fine mesh (1-2mm) stainless steel gauze to support the sorbent
- Course mesh zinc coated steel grate (30-60 mm) to support gauze
- Sides (1m high)
- Cover, corrugated roofing sheets (bitumen)
- Steel structure to support the grate/sides/roof
- Main valve
- Filling system
Fine mesh (1-2mm) stainless steel gauze to support the sorbent
1545.34 m2 € 3.41 €/m2 € 5,263.87
Course mesh zinc coated steel grate (30-60mm) to support gauze
1545.34 m2 € 1.50 €/kg € 2,318.01
Sides (1m high) 7525.78 kg € 1.35 €/kg € 10,159.81
Cover, corrugated roofing sheets (bitumen) 2231.38 m2 € 4.00 €/m2 € 8,925.51
Steel structure to support the grate/sides/roof 1967.59 m € 50.00 €/m € 98,379.31
6.2 Estimation of CAPEX of hydrogen production by electrolysis
Hydrogen electrolyser
Not many manufacturers of electrolyzers are active around the world. The simple reason is that for
nearly all applications the cost of hydrogen is lower if bought directly in bottles, skids, tanks or
pipelines and for high volume consumption applications production is cheapest from natural gas. So,
the market for electrolyzers is limited.
17
Datasheet EU 1000 - EU 1000/B from Agilent technologies (REFC08)
antecy 23
Hgenerators.com has a list on their website with prices for various sizes of electrolyzers (REFH01):
Adjustable Rates / Model
H2O consumed ml/hour
kWh consumed per hour
H2 gas produced
ml/min
H2 gas produced liters/hour
H2 gas produced cubic meters/hour
Affordable Price:
LM-200 10 0.1 200 12 0.012 $4,995
LM-300 15 0.15 300 18 0.018 $5,995
LM-500 25 0.25 500 30 0.03 $6,595
LM-1000 50 0.5 1000 60 0.06 $9,895
LM-2000 100 1 2000 120 0.12 $14,295
LM-3000 150 1.5 3000 180 0.18 $17,595
LM-5000 250 2.5 5000 300 0.3 $23,095
LM-10000 500 5 10000 600 0.6 $42,895
LM-20000 1000 10 20000 1200 1.2 $85,795
LM-30000 1500 15 30000 1800 1.8 $128,685
LM-60000 3000 30 60000 3600 3.6 $257,995
HG-50 41667 215 833333 50000 50 $398,995
HG-100 83334 430 1.6MM 100000 100 $598,995
HG-200 166668 860 3.3MM 200000 200 $798,995
From these prices a price per m3H2/hr capacity can be calculated. These values can be plotted as
datapoints to find the relation between size and price for electrolyzers.
It can be seen that the price per m3H2/hr converges to around 4000 USD per m3/hr installed for large
units. The base calculation case uses roughly 900 m3/hr, which is thus estimated at 3.6 mlnUSD
(2.67 mln€). It can be expected that the cost for hydrogen can further decrease in the future, when
larger installations are built.
Hydrogen electrolyzer, calculation method #2
Since hydrogen production by electrolyzer is not a technology that is widely used: for small scale
applications hydrogen procurement in bottles or trucks is generally more economical, for large scale
applications the same applies for production from natural gas. The Chlor-alkali process on the other
y = 56216x-0.443
0
50000
100000
150000
200000
250000
300000
350000
400000
450000
0.01 0.1 1 10 100 1000
Pri
ce (
USD
) p
er
m3/h
r
Unit size (m3/hr)
Reeks1
Fit (power)
antecy 24
hand, is widespread and applied on kton/year scale all over the world. The electrolyzers used in this
process are not fundamentally different from hydrogen electrolyzers, thus calculating the cost of
such an electrolyzer for hydrogen production may be insightful for large scale hydrogen production.
From18 it is found that for a 160.000 Mton/year chlor-alkali plant, the cost breakdown is as follows:
Item Estimated investment in kUS$
Cells 27,200
Brine purification 14,000
Chlorine processing 16,000
Waste gas treatment 2,300
Caustic evaporation 6,900
Utilities 4,500
Rectifiers 10,000
Engineering 10,000
Total 90,900
Thus, the cost of the electrolysis cells is 27,2 mln USD. The hydrogen production of a
160.000 Mton/yr chlorine plant is 55 mln m3H2/yr or 6,3·103 m3H2/hr. This brings the cost per m3/hr
installed to 4334 USD per m3/hr installed, nearly identical to the previous analysis.
A further conclusion of this analysis is that, at the moment, economics of production scale do not
apply to electrolyzers. This was included in the scale up calculation.
6.3 Estimation of CAPEX of methanol synthesis
Hydrogen compressor
Most commercial hydrogen electrolyzers produce hydrogen at elevated pressure, but rarely above
10barg. For this reason a compressor is needed to compress the produced hydrogen further to at
least 50 barg. To estimate the cost of a the hydrogen compressor a quote from RixIndustries was
obtained. This compressor is rated at 750 SQFM (1275 m3/hr), with suction at 110 psig (7.5 barg) and
discharge at 900 psig (62 barg). This is somewhat higher than needed (required is approx. 1000 m3/hr
at 50 bar). The price quoted is 375 kUSD, or 280 k€. This is not likely to be the cheapest or best
option, but will serve for calculation and estimation purposes.
CO2 compressor
In principle the compressor selected for calculation for compression of hydrogen is large enough to
compress both the CO2 and hydrogen at the same time. However, the CO2 is produced at
atmospheric pressure, so the pressure needs either to be lifted to 7,5 barg when mixing with
hydrogen is done before the hydrogen compressor, or compressed to 50 barg when mixing with
hydrogen is done after the hydrogen compressor. In either way, the flow of CO2 is one third of the
hydrogen flow, and CO2 is an easier gas to compress (both mechanically in terms of leak tightness
and combustibility as well as thermodynamically). For this reason the cost for the additional CO2
compressor is estimated at just over one third of the cost of the hydrogen compressor, at 100k€.
18
Supramaniam Srinivasan, “Fuel cells: From Fundamentals to Applications”, ISBN-10: 0-387-25116-2, 2006 Springer, table 3,7 on page 114
antecy 25
Methanol reactor
The methanol reactor will be built under license modelled on known and tested blueprints. To
estimate the costs of this part, the George Olah methanol plant in Iceland is used as a reference. The
technology used in this plant is almost identical to what is being developed by ANTECY (with the
main exception that ANTECY has the capability to obtain CO2 from air). The size of the Iceland plant is
also neigh on the same as the base calculation case of ANTECY. According to literature (REFM02) this
plant was developed at a cost of 6mln€ (8 mln USD). Applying the reverse calculation for cost for
Connections, Engineering and Startup, this price would imply that the total cost of the plant’s
hardware was 4mln€. Within this 4mln there is the cost for hydrogen electrolyzers (3 mln€),
hydrogen compressors (275 k€), CO2 compressors (100 k€), leaving 600 k€ for the methanol reactor
& distiller.
Alternatively one could look at the cost of a large methanol plant. Off course the technology for this
type of plant is somewhat different from what is being proposed by ANTECY (conventional methanol
production is done from methane/NG). According to REFM01, a 5000 ton/day methanol plant costs
480 mln USD to build. Again removing the cost for Connections, Engineering and Startup, this leaves
a hardware total of 320 mln USD. A plant of this size is roughly 550 times bigger than the base
calculation case for ANTECY. Dividing the cost of 320 mln USD by 550 directly gives a cost for a
methanol plant of 580 kUSD or 435 k€, including the hydrogen compressors and CO2 compressors.
This means that the cost estimated through the George Olah comparison of 1 mln€ is significantly
higher. However, a straight division is not likely to give an accurate result since there is both a
difference in technology as well as a significant difference in size (which leads to economics of scale
differences). In spite of this, the calculation does indicate that the price estimate through the
comparison with the George Olah plant yields a credible result.
6.4 Total main equipment cost Of the costs estimated, the hydrogen electrolyzer constitutes the bulk of the costs. The new concept
(multiple large adsorbers coupled to two desorbers) reduces the cost main equipment of the CO2
capture by roughly 40% (from 2,85 mln€ to 1,7 mln€). It is also expected to greatly reduce the cost of
valves, piping and electronics/data acquisition and control hardware and.
Item Cost (k€)
CO2 adsorption
Desorbers (2) 86
Vacuum pump 50
Adsorbers (5) 125
Sorbent 628
Blower 300
Total CO2 adsorption 1189
H2 generation
Electrolyzer 2667
Total H2 generation 2667
Methanol synthesis
Hydrogen compressor 280
CO2 compressor 100
Reactor 600
Total Methanol synthesis 980
Total hardware 4836
antecy 26
6.5 Real construction cost The cost of main equipment for the 5 MW plant has been estimated at around 5mln€. However, for
many parts, additional cost is required for installation of the equipment. This is only not the case for
the electrolyzer (which is bought as a complete “plug-and-play” unit) and the sorbent (which is just
inventory). For the other parts, a rule of thumb of a factor four is used to install the equipment. This
brings the total battery limit cost of the plant to 9459 k€.
Over the total battery limit cost, a 50% increase is calculated to account for utilities. In the case of
the 5 MW unit, this amounts to 4730 k€. This brings the total cost estimate to 14189 k€.
This is summarized in the following table:
Main equipment cost (k€) Multiplication Battery limit cost (k€)
H2 generator 2667 1.00 2667
H2 compressor 280 4.00 1120
CO2 compressor 100 4.00 400
Reactor 600 4.00 2400
Vacuum pump 50 4.00 200
Blower 300 4.00 1200
CO2 capture unit 125 4.00 500
Desorber tank 86 4.00 344
Sorbent 628 1.00 628
Battery limit total 9459
Utilities 0.5 4730
Total 14189
6.6 Discounting cost To get a realistic ideal for the economic feasibility of the concept, the Net Present Value is calculated
over a 25 year period, with various discounting rates. The following assumptions are made:
The investment of the construction of the plant is done over two years: two thirds of the
investment is made in year one, one third in year two.
Production starts in year two, with half production that year. Full production is assumed for
years three and beyond.
The plant runs year round, on average, at 80% capacity.
Operational cost (excluding energy cost) is estimated at 100 k€ per year for year three and
beyond, with a 2% rise per year to account for inflation.
Energy costs are assumed constant at 0,02 €/kWh.
Methanol value is assumed to be 397 €/ton in year one, and increased by 5% per year to
account for inflation and increased scarcity of fossil raw materials (the validity of this
assumption is verified in Figure 10, the 5% per year increase line is shown in red).
antecy 27
Figure 10: The historical evolution of methanol price
The result can be used to calculate a cost of production of methanol through the solar fuels process,
and compare this to the projected methanol price. The calculation is done by dividing the discounted
kg’s produced over the 25 year period by the discounted total cost over the same period. In the
following table, it can be seen that the discounted cost per discounted kg are 0.66 €/kg and 0.86 €/kg
respectively for 5% and 10% discount rates. So, it can be concluded that, the 5 MW plant would not
be competitive with current methanol sources (but this would not be the purpose of the 5 MW
plant).
€ 0
€ 100
€ 200
€ 300
€ 400
€ 500
€ 600
apr-01 jan-04 okt-06 jul-09 apr-12 dec-14
His
tori
cal M
eth
ano
l pri
ce (
€/t
on
)
Date
antecy 28
Year Capital cost Fixed operational cost Energy cost CO2 tax Total cost Production (kg) 1 € 9.47 M € 50.0 k € .0 k € - € 9.52 M 0 2 € 4.73 M € 50.0 k € 359.93 k € - € 5.14 M 1425693 3 € - € 100.0 k € 719.86 k € - € 819.86 k 2851386 4 € - € 102.0 k € 719.86 k € - € 821.86 k 2851386 5 € - € 104.04 k € 719.86 k € - € 823.90 k 2851386 6 € - € 106.12 k € 719.86 k € - € 825.98 k 2851386 7 € - € 108.24 k € 719.86 k € - € 828.10 k 2851386 8 € - € 110.41 k € 719.86 k € - € 830.27 k 2851386 9 € - € 112.62 k € 719.86 k € - € 832.48 k 2851386 10 € - € 114.87 k € 719.86 k € - € 834.73 k 2851386 11 € - € 117.17 k € 719.86 k € - € 837.03 k 2851386 12 € - € 119.51 k € 719.86 k € - € 839.37 k 2851386 13 € - € 121.90 k € 719.86 k € - € 841.76 k 2851386 14 € - € 124.34 k € 719.86 k € - € 844.20 k 2851386 15 € - € 126.82 k € 719.86 k € - € 846.69 k 2851386 16 € - € 129.36 k € 719.86 k € - € 849.22 k 2851386 17 € - € 131.95 k € 719.86 k € - € 851.81 k 2851386 18 € - € 134.59 k € 719.86 k € - € 854.45 k 2851386 19 € - € 137.28 k € 719.86 k € - € 857.14 k 2851386 20 € - € 140.02 k € 719.86 k € - € 859.89 k 2851386 21 € - € 142.82 k € 719.86 k € - € 862.69 k 2851386 22 € - € 145.68 k € 719.86 k € - € 865.54 k 2851386 23 € - € 148.59 k € 719.86 k € - € 868.46 k 2851386 24 € - € 151.57 k € 719.86 k € - € 871.43 k 2851386 25 € - € 154.60 k € 719.86 k € - € 874.46 k 2851386 Discounted costs per discouted kg Sum € 14.20 M € 2.98 M € 16.92 M € 34.10 M 67007581.54 € 0.509 Discount rate 5% € 13.97 M € 1.64 M € 9.59 M € 25.21 M 37987457.15 € 0.664 Discount rate 10% € 13.77 M € 1.03 M € 6.14 M € 20.94 M 24322892.59 € 0.861 Discount rate 15% € 13.58 M € 720.0 k € 4.32 M € 18.62 M 17105435.48 € 1.089
5 MW: base case
antecy 29
Appendix II: detailed cost estimate of the 500 MW case
In order to evaluate the effect of scale up on the economics of the process, rules of thumb were used
to calculate the cost of the process upscaled to 500 MW. In order to come to a realistic rule of
thumb, the data in a cost estimation article19 of the cost of CO2 sequestration by an amine system is
used as a reference. The article calculates a 600 MEe coal power plant. At 820 gCO2/kWh, 7500 hours
per year and 90% capture rate, this is 3,32e6 ton CO2/year. Our 5 MW case sequestrates 3,8e3
ton/year, so our 5 MW case is approximately 850 times smaller.
A flat multiplication of two comparable parts: the blower and the CO2 compressor show that this
highly overestimates the costs. The 5 MW case CO2 compressor investment cost, 400 k€, multiplied
by 850, gives 340 mln€. Cost cited in19 is 31,73 mln€. The 5 MW case blower is estimated at 30 k€ (for
the flue gas case), times 850 gives 25,5 mln€. The cited cost is 3,1 mln€.
Alternatively, one could use a multiplication to the power 0.66. This gives, both for the blower and
compressor a much more accurate result (34,3mln€ and 2.6mln€ respectively). So, for these parts, as
well as the H2 compressor, methanol reactor and vacuum pump, this calculation is used. For the
electrolyzer the benefit of scale up is expected to be less pronounced, so a cost reduction of 10% is
applied (cost multiplied by 90 instead of 100). For the cost of desorption vessels, adsorption vessels
and sorbent a flat multiplication is used. This yields the following result:
5 MW Battery limit cost (k€)
Scale up 500 MW cost estimate (k€)
H2 generator 2667 x90 240030
H2 compressor 1120 A*x^0,66 23400
CO2 compressor 400 A*x^0,66 8357
Reactor 2400 A*x^0,66 50143
Vacuum pump 200 A*x^0,66 4179
Blower 1200 A*x^0,66 25072
CO2 capture unit 500 x100 50000
Desorber tank 344 x100 34400
Sorbent 628 x100 62800
Battery limit total 9459
498381
Utilities 4730 249191
Total 14189
747572
Using this cost estimate, the cost of producing methanol can be estimated, similarly to was done for
the 5 MW case. The discounted cost per discounted kg are 0.489 €/kg and 0.60 €/kg respectively for
5% and 10% discount rates.
19
Abu-Zahra et.al. Int.Journ. of Greenhouse Gas Control I, (2007) 135-142
antecy 30
Year Capital cost Fixed operational cost Energy cost CO2 tax Total cost Production (kg) 1 € 498.67 M € 500.0 k € - € - € 499.17 M 0 2 € 249.33 M € 500.0 k € 35.99 M € - € 285.83 M 142569322 3 € - € 1.0 M € 71.99 M € - € 72.99 M 285138645 4 € - € 1.02 M € 71.99 M € - € 73.01 M 285138645 5 € - € 1.04 M € 71.99 M € - € 73.03 M 285138645 6 € - € 1.06 M € 71.99 M € - € 73.05 M 285138645 7 € - € 1.08 M € 71.99 M € - € 73.07 M 285138645 8 € - € 1.10 M € 71.99 M € - € 73.09 M 285138645 9 € - € 1.13 M € 71.99 M € - € 73.11 M 285138645 10 € - € 1.15 M € 71.99 M € - € 73.13 M 285138645 11 € - € 1.17 M € 71.99 M € - € 73.16 M 285138645 12 € - € 1.20 M € 71.99 M € - € 73.18 M 285138645 13 € - € 1.22 M € 71.99 M € - € 73.21 M 285138645 14 € - € 1.24 M € 71.99 M € - € 73.23 M 285138645 15 € - € 1.27 M € 71.99 M € - € 73.25 M 285138645 16 € - € 1.29 M € 71.99 M € - € 73.28 M 285138645 17 € - € 1.32 M € 71.99 M € - € 73.31 M 285138645 18 € - € 1.35 M € 71.99 M € - € 73.33 M 285138645 19 € - € 1.37 M € 71.99 M € - € 73.36 M 285138645 20 € - € 1.40 M € 71.99 M € - € 73.39 M 285138645 21 € - € 1.43 M € 71.99 M € - € 73.41 M 285138645 22 € - € 1.46 M € 71.99 M € - € 73.44 M 285138645 23 € - € 1.49 M € 71.99 M € - € 73.47 M 285138645 24 € - € 1.52 M € 71.99 M € - € 73.50 M 285138645 25 € - € 1.55 M € 71.99 M € - € 73.53 M 285138645 Discounted costs per discouted kg Sum € 748.0 M € 29.84 M € 1691.68 M € 2469.52 M 6700758154 € 0.369 Discount rate 5% € 736.13 M € 16.42 M € 959.03 M € 1711.58 M 3798745715 € 0.451 Discount rate 10% € 725.33 M € 10.32 M € 614.06 M € 1349.71 M 2432289259 € 0.555 Discount rate 15% € 715.48 M € 7.20 M € 431.84 M € 1154.52 M 1710543548 € 0.675
500 MW: base case
antecy 31
Year Capital cost Fixed operational cost Energy cost CO2 tax Total cost Production (kg) 1 € 378.67 M € 500.0 k € - € - € 379.17 M 0 2 € 189.33 M € 500.0 k € 35.99 M € 18.86 M € 206.96 M 171493956 3 € - € 1.0 M € 71.99 M € 37.73 M € 35.26 M 342987912 4 € - € 1.02 M € 71.99 M € 37.73 M € 35.28 M 342987912 5 € - € 1.04 M € 71.99 M € 37.73 M € 35.30 M 342987912 6 € - € 1.06 M € 71.99 M € 37.73 M € 35.32 M 342987912 7 € - € 1.08 M € 71.99 M € 37.73 M € 35.34 M 342987912 8 € - € 1.10 M € 71.99 M € 37.73 M € 35.36 M 342987912 9 € - € 1.13 M € 71.99 M € 37.73 M € 35.38 M 342987912 10 € - € 1.15 M € 71.99 M € 37.73 M € 35.41 M 342987912 11 € - € 1.17 M € 71.99 M € 37.73 M € 35.43 M 342987912 12 € - € 1.20 M € 71.99 M € 37.73 M € 35.45 M 342987912 13 € - € 1.22 M € 71.99 M € 37.73 M € 35.48 M 342987912 14 € - € 1.24 M € 71.99 M € 37.73 M € 35.50 M 342987912 15 € - € 1.27 M € 71.99 M € 37.73 M € 35.53 M 342987912 16 € - € 1.29 M € 71.99 M € 37.73 M € 35.55 M 342987912 17 € - € 1.32 M € 71.99 M € 37.73 M € 35.58 M 342987912 18 € - € 1.35 M € 71.99 M € 37.73 M € 35.60 M 342987912 19 € - € 1.37 M € 71.99 M € 37.73 M € 35.63 M 342987912 20 € - € 1.40 M € 71.99 M € 37.73 M € 35.66 M 342987912 21 € - € 1.43 M € 71.99 M € 37.73 M € 35.69 M 342987912 22 € - € 1.46 M € 71.99 M € 37.73 M € 35.71 M 342987912 23 € - € 1.49 M € 71.99 M € 37.73 M € 35.74 M 342987912 24 € - € 1.52 M € 71.99 M € 37.73 M € 35.77 M 342987912 25 € - € 1.55 M € 71.99 M € 37.73 M € 35.80 M 342987912 Discounted costs per discouted kg Sum € 568.0 M € 29.84 M € 1691.68 M € 1402.90 M 8060215939 € 0.174 Discount rate 5% € 558.98 M € 16.42 M € 959.03 M € 1031.80 M 4569439764 € 0.226 Discount rate 10% € 550.79 M € 10.32 M € 614.06 M € 853.33 M 2925754997 € 0.292 Discount rate 15% € 543.30 M € 7.20 M € 431.84 M € 756.01 M 2057580658 € 0.367
500 MW: most optimistic
antecy 32
Appendix III: projected cost of PV and green hydrogen
Figure 11. Base case and aggressive scenarios for cost projections of solar PV power, electrolyser CAPEX, and hydrogen.
Figure 11 shows the projected price of PV electricity, electrolyser CAPEX and green hydrogen. The
following assumptions were used:
Net present value is assumed zero at a discount rate of 7%, equivalent to a weighted average
cost of capital of 7%.
32% utilisation of electrolyser.
Electrolyser capacity of 75% of PV DC capacity (to increase electrolyser utilisation).
A solar yield of 23.7% for single axis tracking system in China (includes 10% curtailment of
electricity because of undersized electrolyser) is used to calculate cost of electricity.
Installed costs of electrolyser 1.2 times electrolyser costs.
5% manufacturing margin.
The data presented is the expectation for modular, MW scale units with minimal installation costs.
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Green H2 costs scenarios Fossil H2 $/kg
LNG based, 12$/mmBTU, LCOE $c/kWh
4-10 $/mm BTU NG, electricity, VarCo $c/kWh
green H2, Base $/kg
green H2, Aggressive $/kg
PV electricity, Base $c/kWh
PV electricity, Aggressive $c/kWh
Electrolyser Capex, Base $/kW
Electrolyser Capex, Aggressive $/kW