Skoglund et al.
10.3389/fther.2023.1282028
TABLE 2 Typical fl ue gas data for the case study mill.
Flue gas fl ow [m 3 /h dry]
Emission source
Temperature [ ° C]
CO
2 conc. [vol% dry]
H 2 O content [vol%]
Lime kilns
45 000
300
22
28
Small recovery boiler
100 000
54
12
13
Large recovery boiler
300 000
170
16
27
Power boiler
110 000
57
8
25
It is further assumed that 90% of the CO 2 fromthe fl ue gases of the two recovery boilers and the lime kilns is captured, with a common absorber treating the mixed fl ue gases from all three stacks. TheCO 2 from the power boiler is not considered for capture, partly because the power boiler can have large operational variations as described above, and partly because the load of the power boiler, and consequently its CO 2 emissions, is highly dependent on the energy balances of the integrated processes at the mill site, and therefore would differ considerably between the different integration scenarios. A simpli fi ed fl owsheet of the carbon capture process using AMP-PZ solvent is shown in Figure 2. The CO 2 -rich fl ue gas feed is cooled to around 40 ° C before being sent to an absorber forCO 2 absorption. A water wash section is placed at the top of the absorber to recover solvents from the CO 2 -lean fl ue gases. The rich solvent is extracted from the absorber bottom and sent to the stripper for regeneration. The CO 2 captured is extracted from the stripper top and sent to the subsequent conditioning unit. The solvent is thereafter regenerated, i.e., lean solvent is extracted from the stripper bottom and sent back to the absorber for reuse. A lean- rich heat exchanger is used for heat recovery from the lean solvent. The carbon capture process was simulated using rate-based models in Aspen Plus V10. The eNRTL property method was used. The reactions involved, the equilibrium and kinetics parameters and interaction parameters were taken from Aspen Plus default library models for AMP and PZ solvents and literature (Li et al., 2014; Zhang et al., 2017). The process was simulated with fl ue gas data representing normal mill operation as speci fi ed in Table 2. High- temperature fl ue gases were assumed to be cooled to 125 ° C before entering the direct contact cooler, in order to enable recovery of as much heat as possible without risking condensation of acids in the heat exchangers. The CO 2 leaving the stripper was further assumed to be conditioned in a liquefaction plant to the conditions required for ship transport, before being sent to a buffer storage. In the liquefaction plant, the CO 2 -rich stream from the capture process is compressed in a multi-stage compressor train with intercoolers, and lique fi ed using an ammonia refrigeration cycle. In the compression train, condensed water is removed from the feed gas through phase separators. Deep water removal is achieved with a dehydration unit. Other impurities, mainly O 2 , are removed as a purge gas after the liquefaction, resulting in a minor loss of the CO 2 . The process also includes a recirculation of fl ash gas that is generated during a throttling step for further puri fi cation of the liqui fi edCO 2 . The liqui fi ed CO 2 has a pressure of 16 bar(a), a temperature of − 26 ° C, and is puri fi ed to 99.999 mol% CO 2 and 0.001 mol% N 2 to ful fi l the Northern Lights speci fi cations (Equinor, 2019). The
Process cooling in the mill is achieved via the secondary heat recovery system. Process streams are cooled with water, which is heated to warm or hot water temperatures and sent to the water tanks of the secondary heating system. The water is then used for heating and as process or cleaning water in various parts of the mill. Additional cooling at the mill is done using cooling towers to cool dirty water before wastewater treatment. As a basis for the heat integration analysis (see Section 2.4), all process heating and cooling demands in the mill were mapped according to their heat fl ows and temperature levels (see Section 12). Table 2 reports fl ue gas data for all CO 2 emission sources in the mill. The two recovery boilers have separate stacks, with different fl ue gas treatment systems installed. Flue gases from the larger recovery boiler go through a convective section with a superheater and economizer before being passed through an electrostatic fi lter and then sent to the stack. In the smaller recovery boiler, the fl ue gases also go through a superheater fi rst, but then pass through an electrostatic fi lter before going to the economizer. These fl ue gases also pass through a scrubber, where they are further cleaned and cooled before being sent to the stack. The hot fl ue gas from the lime kilns goes through an electrostatic fi lter before being sent to the stack. While the recovery boilers and lime kilns are operated based on demand for recovering chemicals from the black liquor, the power boiler is operated to balance total steam production with the process steam demands. Consequently, the load of the power boiler varies quite signi fi cantly, and will also differ between the integration scenarios. The numbers in the table are only meant to provide an indication of typical current emission levels from the boiler, which are not considered for capture. In order to estimate the change in carbon fl ows related to the power boiler, the CO 2 emission factor of the solid wood fuel used in the power boiler was assumed to be 96 g/MJ (Naturvårdsverket, 2005) and the fuel to heat ef fi ciency of the boiler was assumed to be 85%.
2.2 Modeling of the carbon capture and liquefaction plant
The carbon capture process assumed in this paper is a post combustion amine absorption process, since this is one of the most mature options for carbon capture from fl ue gases, which is also well-suited for retro fi t of existing industrial plants such as pulp and paper mills. The capture process was assumed to use the AMP (2- amino-2-methyl-propanol) -PZ (piperazine) solvent that has been suggested as the new benchmark solvent for post-combustion carbon capture (Feron et al., 2020).
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