Aspen
HYSYS simulation and economic analysis of CO2 removal from flue gas by amine
absorption for CCS and EOR
Abstract
This was an MSc
project undertaken in order to study the feasibility of a post combustion
carbon capture plant using amine absorption if a carbon tax £/tonne CO2 were
introduced. Aspentech HYSYS software was used to model the carbon capture plant
whereby carbon dioxide (CO2) is removed using a Monoethanolamine (MEA) solution
by absorption from a flue gas exiting a combined cycle gas turbine (CCGT)
plant. This study has incorporated two configurations to model the CO2 capture process,
a base case configuration and a vapour recompression configuration. The base
case model achieved a CO2 removal of 86% and heat consumption in the reboiler
of 3.48 MJ/kg CO2 removed. Its total capital cost was £110.63 million and an
net present value (NPV) of -£490.29 million over 20 years at a 7% discount
rate. The vapour recompression model achieved a CO2 removal of 88% and heat
consumption in the reboiler of 3.19 MJ/kg CO2 removed. Its total capital cost
was £160.67 million and an NPV of -£551.94 million. An MEA concentration of
32-33% for the lean amine stream entering the absorber column was been found
give the lowest heat consumption value in the reboiler. It has also been shown
that the first stage is ideal for the rich amine to enter the desorber with the
vapour stream entering directly underneath. A flash drum pressure of 110 kPa
produced the lowest operating costs however, it was not known if there are more
optimal pressures as the model failed to converge for higher pressures. The
findings produced a minimum carbon tax of 31.01 and 34.04 £/tonne CO2 for the
base case and the vapour recompression model respectively in order to achieve
an NPV of 0. As sale of CO2 is in £/tonne, the price at which the CO2 is sold
for can be subtracted directly from the minimum carbon tax value needed in
order to obtain a new minimum that can achieve an NPV of zero.
Introduction
Increasing
concentrations of greenhouse gases such as CO2 into the atmosphere is one of
the leading causes of global warming, therefore governments and institutions
are actively looking for a way reduce emissions of CO2. A significant amount of
CO2 emissions come from power plants used to generate electricity from the
combustion of fossil fuels such as coal, natural gas and other CO2 emitting
sources (Committee on climate change, 2015). However, with world energy demand
increasing yearly and current constraints on non-greenhouse sources of energy
such as cost, availability and storage along with the time needed to
transition, it is likely that fossil fuels are still the main source of energy
for many years to come (EIA, 2016). Removal of carbon dioxide using amine
absorption from streams is not a new concept in industry however, it is
generally applied to high pressure streams from oil field reservoirs. The
reason for this is to increase heating value, lessen the load on the
compressors and minimise potential corrosion of pipes. Post combustion removal
from electricity generation facilities has not been done on a large scale due
to the likely cost restrictions it presents (Caldecott et al., 2016). Storage
of captured CO2 also presents major problems as it incurs the additional cost
of transportation and pumping into geological formations. There is also the
risk of the leakage of CO2 back into the atmosphere.
CO2 can be
injected into an oil field for enhanced oil recovery (EOR) in order to recover
more oil than would be feasible using traditional methods therefore, an ideal
scenario would be to sell the captured CO2 to oil companies for EOR although
that option is not always readily available. The most common capture
technologies employ amine gas treating which involves the CO2 being absorbed
into the amine solution in an absorber column. The amine is then regenerated
from the resultant stream in a desorber column by the reboiler and stripper.
The amine is recycled back into the absorber and the CO2 stream is condensed to
remove water before it is ready to store. Diethanolamine (DEA),
Monoethanolamine (MEA), Methyldiethanolamine (MDEA) are the most commonly used
amines for gas treating.
A carbon capture
pilot plant testing facility TCM exists in Mongstad, Norway (MIT, 2012). This
pilot plant is being used to test two alternative post combustion CO2 removal
techniques, one of which is using chilled ammonia and the other is using
amines. It is a joint venture being undertaken by the Norwegian state, Statoil,
Shell and Sasol. The project was undertaken in order to use the information
gathered to build a full scale plant however, it was decided in September 2013
by the Norwegian Oil and Energy Ministry that it will drop plans to build a
full scale plant. Regardless, the pilot plant has remained operational as a
testing facility.
Aspentech HYSYS simulation software has been used to
model two plant configurations as shown in Figures 1 and 2 (Birkelund, 2013).
Figure 1: Base case configuration
Figure 2: Vapour recompression configuration
Base case
The amine fluid
package in HYSYS was used which comes pre-prepared with the relevant reactions
and is optimised for amine solvents in general.
The model
consists of a flue gas fan, an absorber column, rich amine pump, rich/lean heat
exchanger, desrober column, condenser, reboiler, lean pump, lean cooler, CO2
cooler and a separator. A direct contact cooler (DCC) is modelled by a flue gas
cooler in order to simplify the simulation. HYSYS also enables the addition of
a make-up object which can help simplify the addition of lost components to the
stream before being recycled. Pressure drops across the various units have not
been taken into account. Hot flue gas
from the power plant enters a transport fan before being cooled to 40˚C
(Kallevik, 2010) in the DCC. The flue then enters the absorber column from the
bottom and a lean amine stream enters from the top. The lean amine stream is a
stream that consists of approximately 29 wt% MEA, 64.5 wt% water and 5.5 wt%
CO2. The CO2 present in the stream is because not all of the CO2 is removed
during the regeneration process. Sweet
gas exits the absorber at the top and is released into the atmosphere.
Additional water is recycled through the absorber column to aid the process. A rich
amine stream exits the bottom of the absorber and is pumped to 200kPa by the
rich amine pump. The pressurised stream enters the lean/rich heat exchanger
where the hot lean amine stream to be recycled heats the rich amine stream to
104-109.5˚C. The rich amine stream then enters the desorber column where
regeneration of the amine takes place. The amine is heated to 120˚C in the
reboiler (Øi, 2007) and the CO2 and some water leave the top of the desorber as
vapour and the lean amine liquid stream exits at the bottom. The recovered CO2
stream is cooled before it enters a separator to split the CO2 from the water.
The lean amine stream feed is then pumped into the lean/rich heat exchanger in
order to reduce its temperature and provide heat for the rich amine stream. The
lean amine stream temperature that exits the lean/rich heat exchanger is still
too high therefore it undergoes further cooling. Any lost MEA or water is added
back before being recycled back into the Absorber.
Vapour
recompression
The vapour
recompression model contains the same units as the base case model with the
addition of a lean amine pump, a flash drum and a compressor. The purpose of
this set up is to reduce the reboiler duty through the addition of a recycle
vapour stream in the desorber (Karimi et al., 2016).
In the vapour recompression
configuration, the lean amine stream exiting the desorber column is reduced to
101 kPa by the pressure reducing valve which has the effect of reducing the
temperature to 102˚C. This stream enters the flash drum which then separates
some of the water from the lean amine stream. The liquid stream (lean amine)
exits at the bottom of the vessel where it is then pumped to 120 kPa. This
stream then enters the lean/rich heat exchanger and is recycled back into the
desorber using the same steps as the base case model. The vapour stream
consisting of 95.5 mol% water vapour exits at the top of the flash drum and
enters the compressor where it is compressed to 200 kPa (desorber column
pressure). The pressure increase on the vapour stream has the effect of
increasing the temperature to 190˚C before it enters the desorber to provide
additional heat to the system.
Results
In order to
calculate the NPV the capital costs and operating costs need to found. The
costs of the major equipment were calculated by applying scaling factors,
currency conversion and price index equations to similar work in carried out by
Vozniuk (2010). Once the costs were determined an estimate of the capital costs
can be made using Lang factors. The Lang factor is a ratio of the total cost of
installing the plant to the cost of the total purchase equipment cost (Sinnott,
2005).
The cost of the
liquefaction plant has been included as an estimate from Øi et al. (2016). The
capital costs are shown in Table 1. The operating costs are calculated for the
two models by calculating the energy consumption per year (8000 hr.) and
multiplying it to the utility cost estimates from a CCGT plant (Department for
Business, Energy & Industrial Strategy, 2016) and are shown in Table 2. The
net NPV is calculated for a period of 20 years at a 7% discount rate.
Table 1: Capital costs
Table 2: Operating costs
Sensitivity
analysis
Case studies
were run in HYSYS in order to see the effect of changing the parameters on NPV
and heat consumption. The sensitivity analysis for NPV incorporated a 10
£/tonne CO2 carbon tax in order to fairly assess variations where operating
cost is higher but the captured CO2 is also higher. NPV has also been taken as
positive therefore, for the sensitivity analysis graphs a lower NPV represents
a lower investment cost than a higher NPV.
Variation of MEA
concentration
The
concentration of MEA was varied from 30-36 wt% in order to see the effects on
heat consumption in the reboiler [MJ/kg CO2] (Figure 3).
Figure 3: Heat consumption [MJ/kg CO2] as a
function of MEA concentration
It can be seen
in figure 3 that an MEA concentration of 32-33% for the lean amine stream
entering the absorber column gives the lowest heat consumption value in the
reboiler. However, the range of values for MEA wt% needs to be increased in
order to accurately assess the findings.
Inlet stage
variation
The inlet stage
of the rich amine stream was varied while keeping the reboiler duty the same in
order to find the heat consumption in the reboiler [MJ/kg CO2] and the reboiler
temperature (Figure 4).
Figure 4: Heat consumption
[MJ/kg CO2] and reboiler temperature (°C) as a function of inlet stage of rich
amine stream into the desorber for vapour recompression model
It can be seen
that having the rich amine stream enter the absorber in the first stage gave
the lowest value for heat consumption. This is because it gave the highest
reboiler temperature for the same duty. It was found that in order to provide
the best heat for the rich amine stream the vapour recycle stream would need to
enter in the stage directly under it.
Variation of
flash drum pressure
A case study on
the effects of changing the flash drum pressure in the vapour recompression
model has been plotted in Figure 5.
Figure
5: Heat
consumption [MJ/kg CO2] as a function of MEA concentration
It can be seen
that as the flash drum pressure increases the heat consumption by the reboiler
(MJ/kg CO2 removed) increases but the NPV decreases. This is because although
there are higher operating costs from the reboiler when the flash drum pressure
is increased the vapour outlet decreases therefore, the compressor operating
costs decrease. The study could not converge beyond 110 kPa therefore it is
unknown if the NPV would decrease more at a higher flash drum pressure and when
the point is reached where the NPV stops decreasing. These findings are similar
to those found by Birkelund (2013) who has analysed the effects of changing the
flash drum pressure on Equivalent work of the whole process [kJ/kg] which is
comparable to the NPV findings for this study as capital costs did not change
with respect to the flash drum pressure. Birkelund (2013) has found that the
equivalent work decreases until approximately 115 kPa where it then starts to
increase.
NPV variations
based on Carbon Tax and income of CO2
In order to
estimate a carbon tax price (£/tonne CO2) that would results in a 0 or positive
NPV, carbon tax was treated as income in the NPV calculations.
Figures 6 and 7
show the change in NPV if the CO2 captured was sold for EOR for the base case
model and vapour recompression model respectively.
Figure 6: NPV as
a function of carbon tax in the range of 0-35 £/tonne CO2 incorporating income
obtained from the sale of CO2 for EOR at 0, 5, 10, and 15 £/tonne CO2 for the
base case configuration
Figure
7: NPV as a
function of carbon tax in the range of 0-35 £/tonne CO2 incorporating income obtained
from the sale of CO2 for EOR at 0, 5, 10, and 15 £/tonne CO2 for the vapour
recompression configuration
The figures show
a linear relationship between carbon tax and sale price to NPV. Goal seek
function was used in order to find the carbon tax needed for NPV to equal 0 for
both the base case model and the vapour recompression model not taking in into
account income received from the sale of
CO2. The resultant carbon tax needed was 31.01 and 34.04 £/tonne CO2 for the
base case and the vapour recompression model respectively.
Calculations
used to find NPV assumed no additional operating or capital costs incurred from
the sale of CO2 based on the assumption of Kemp and Kasim (2012) that with
slight modification most of the pipeline in the UK Central North Sea can be
re-used as they are metallurgically suitable. It can be seen that in order to
have a 0 NPV when including the sale of CO2 £/tonne, the price at which the CO2
is being sold can be subtracted from the minimum carbon tax needed of 31.01 and
34.04 £/tonne CO2 for base case and vapour recompression model respectively to
find the new minimum price of carbon tax for 0 NPV. It is worth noting that
Reid (2015) has that said offshore oil companies in the North Sea may be more
likely to accept the CO2 for free as an exchange for providing free storage for
the captured CO2.
Discussion
The base case
model and vapour recompression model achieved a CO2 removal of 86% and 88% and
heat consumption in the reboiler of 3.48 and 3.19 MJ/kg CO2 removed
respectively. The value achieved for the base case is close to that observed in
literature for Øi (2007) and Karimi et al. (2016). However, the value obtained
for the vapour recompression model is higher than that found in literature
which was 2.7 MJ/kg CO2 removed for Birkelund (2013) and 2.58 MJ/kg CO2 removed
for Karimi et al. (2016). It was seen during simulation that the CO2 present in
the recycle stream has a great effect on the overall CO2 removal % therefore it
is preferable to keep the concentration of CO2 in the recycle stream low.
The main
advantage of the vapour recompression model is that although the additional
units incur a higher capital cost, the reduction in heat consumption of the
reboiler should lower the operating costs enough to make it a viable
alternative, however, it was found that the operating costs saved in the
reboiler was £1.26 million in comparison to the base case model but an
additional cost of £2.2 million was needed for the operating costs of the
compressor therefore, irrespective of the capital costs the base case model
presents the lowest investment cost.
Rich amine
stream entering the absorber in the first inlet stage was found to have the
lowest heat consumption value with the vapour stream entering directly
underneath in order to best provide heat to the system. An optimum MEA
concentration of 32-33% for the lean amine stream entering the absorber column
has been found although the differences on heat consumption were minimal for
the ranges tested. A flash drum pressure of 110 kPa produced the lowest
operating costs but due to the model failing to converge at higher pressures it
cannot be deduced that 110 kPa is the optimal flash drum pressure in the vapour
recompression model and further research is needed.
The minimum
carbon tax £/tonne CO2 needed to achieve an NPV of zero was 31.01 and 34.04
£/tonne CO2 for the base case and vapour recompression model respectively.
These findings show that the proposed carbon tax plans in Canada (The Canadian
Press, 2016) for a minimum of 10 Canadian dollars (£6.09) per tonne of CO2 are
too low to enable a high adoption of carbon capture and storage by amine
absorption. As the CO2 sold for EOR is in £/tonne the selling price can be
deducted directly from the minimum price of carbon tax needed to achieve an NPV
of zero to obtain a new minimum with respect to the sale of CO2. A more precise
costing technique would need to be used in order to verify these results as
Lang factors used in the estimation present a ±50% degree of accuracy.
Recommendations
• Employ the use of a Modified Hysim
Inside-Out algorithm with adaptive damping to improve convergence (Øi, 2007).
• Use a programme to find capital costs as
there are many variables that need to be taken into consideration.
• Research alternate forms of carbon
capture technologies.
Conclusion
The study has
found that carbon tax needs to be in excess of 30 £/tonne of CO2 in order to
motivate companies to set up a carbon capture process as it would be cheaper
for the company to pay the tax then employ CCS for lower carbon tax rates. The
studies have found optimisations to the process for inlet stage and MEA wt%.
There were issues in the convergence of the simulation which should be
addressed for future studies using the technique proposed in the
recommendations. In general the results obtained matched the results found in
literature and therefore, this study with respect to carbon tax and income from
CO2 can be used in order to assess feasibility of similar carbon capture
plants.
Author : Nihad Kassem
Reservoir Engineer
Primera Reservoir LTD
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Strong focus on decision making and uncertainty analysis combined with advanced, state of the art and complex workflows.