Figure 7a shows the cost of capture as a function of the recovery of the
PSA unit for the two MOF sorbents. Cost as a function of other PSA
performance parameters (productivity, purity) are shown in Figure S17.
There is a clear positive correlation between the recovery of the system
and the cost of the CO2 capture system. As the recovery
goes up so too does the size of the downstream liquefaction and recycle
systems. This effect is primarily driven by the amount of gas which has
to be compressed during liquefaction but it is also impacted by the
CO2 purity from the liquefaction column since higher
recoveries are often associated with lower purities.
Figure 7b shows the cost ($/tonneCO2) comparison for
the five MIL-101(Cr) sorbent cases, with each case’s cost broken down
into its constituent costs (capital and energy). As the recovery
increases and the purity decreases, the energy, rotating equipment, and
PSA units’ costs all increase significantly. The increase in energy and
rotating equipment cost mainly comes from the increased amount of gas in
the liquefaction process, which drives up the size and energy demand of
COMP3. In case M5, the energy demand of that single compressor is
greater than the amount of energy produced by the base case plant. The
increase in PSA cost relates both to recovery and productivity. With
increasing recovery more CO2 must be captured by the
PSA, along with slight increases in the amount of CO2
held up in the recycling of the system. Cases M4 and M5 also suffer
from 20-30% reductions in PSA productivity as compared to the lower
recovery cases, which will also play a role in the increased price of
the PSA unit. A similar analysis on UiO-66 sorbent cases is provided in
Figure S18.
As a typical case, the simplified flowsheet for Case M1 is given in
Figure 8. The flue gas enters the system and passes through the direct
contact chiller, where 97.4% of water is removed, reducing the
concentration of water in the stream to 0.47%. The partially
dehumidified gas is then passed through a series of two compressors with
intercoolers, where it reaches a target pressure of 20.49 bar. The
stream is further cooled down to 2 °C in the sub-ambient heat exchanger
network, where additional water is also removed. This high pressure feed
(1) at 20.29 bar contains only 437.6 ppm H2O, which is
removed in an adsorbent dryer. Leaving the adsorbent dryer, the now bone
dry flue gas with composition 14.65% CO2 and 85.35%
N2 is mixed with the cold recycle (6) which is 81.5%
CO2. This mix (2), a CO2 enriched feed
of 18.9% CO2, is then cooled to the -30°C operating
condition of the PSA, entering the system at 20.19 bar.
The N2-rich product in Figure 8 (stream 3), still at a
high pressure condition of 19.99 bar, is then cooled and expanded
repeatedly, generating streams at temperatures as low as -93 °C, while
also recovering 97 MW of energy. The CO2 enriched
product is recovered at 0.103 bar vacuum pressure with a concentration
of CO2 of 93.4%. From the vacuum pump it is
recompressed to a moderate 12.5 bar pressure before being sent back to
the sub-ambient heat exchange network as the feed to the liquefaction
column. Operating at -40.4 °C, the top stream from the liquefaction
column is recompressed from the 12.5 bar liquefaction pressure to 20.29
bar where the cold recycle (stream 6) enriches the feed. The liquid
CO2 product (stream 7), now 99.6% CO2,
is pumped to the feed condition of the 5th stage of
the multistage compression (63.4 bar) and heated to provide heat removal
to the rest of the system from -38 °C to 35 °C. It can then be
compressed to the pipeline condition of 152.7 bar, where it can be
transported to storage or sequestration.