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 CO_{2 ­}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 5^th 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.