The capture of COfrom ambient air was commercialized in the 1950s as a pre-treatment for cryogenic air separation. In the 1960s, capture of COfrom air was considered as a feedstock for production of hydrocarbon fuels using mobile nuclear power plants.In the 1990s, Klaus Lackner explored the large-scale capture of COas a tool for managing climate risk,now commonly referred to as direct air capture (DAC).

Estimates of the cost of DAC vary widely. Cost estimates based on simple scaling relationships yield resultsfrom 50 to 1,000 $/tCO. Uncertainty might be reduced if detailed specifications of individual DAC technologies were available. Yet, despite growing interest in carbon removal as a component of climate strategy, one thorough review,many papers on DAC-to-CCS (carbon capture and storage) comparison,specific absorbers,or components of plausible DAC systems,no prior paper provides a design and engineering cost basis for a complete DAC system for which all major components are (1) drawn from well-established commercial engineering heritage, or (2) described in sufficient detail to allow assessment by third parties. This paper aims to fill that gap.

Plausible DAC processesuse solid sorbentsor aqueous basic solutionsas the capture media. Solid sorbents offer the possibility of low energy input, low operating costs, and applicability across a wide range of scales. The challenges of solid sorbent designs are first, the need to build a very large structure at low cost while allowing the entire structure to be periodically sealed from the ambient air during the regeneration step when temperature, pressure, or humidity must be cycled. And second, the inherently conflicting demands of high sorbent performance, low cost, and long economic life in impure ambient air.

Aqueous sorbents offer the advantage that the contactor can operate continuously, can be built using cheap cooling-tower hardware, and the (liquid) surface is continuously renewed allowing very long contactor lifetimes despite dust and atmospheric contaminants. Once captured, CO 2 can be easily pumped to a central regeneration facility allowing economies of scale and avoiding the need to cycle conditions in the inherently large air contactor. Disadvantages of aqueous systems include the cost and complexity of the regeneration system and water loss in dry environments.

Carbon Engineering (CE) has been developing an aqueous DAC system since 2009. In 2012, we described our air-liquid contactor,the front end of the process. Here, in the next section, we provide an end-to-end overview of our baseline DAC system, proceeding from a high-level description of and heat and mass balance down to descriptions of individual unit operations. The following section provides results from a 1 t-CO/day pilot plant operated since 2015. CE's capital cost estimating process is described in the section on Process Economics along with the levelized cost of capture under various plant configurations and economic assumptions. Finally, the Discussion provides comparison with prior literature and a discussion of options for improving the technology.

Process Description

2 from the atmosphere using an aqueous solution with ionic concentrations of roughly 1.0 M OH−, 0.5 M CO 3 2−, and 2.0 M K+. In the second loop, CO 3 2− is precipitated by reaction with Ca2+ to form CaCO 3 while the Ca2+ is replenished by dissolution of Ca(OH) 2 . The CaCO 3 is calcined to liberate CO 2 producing CaO, which is hydrated or “slaked” to produce Ca(OH) 2 . Figure 1 Process Chemistry and Thermodynamics Show full caption A calcium loop (right) drives the removal of carbonate ion and thus the regeneration of the alkali capture fluid (left). Boxes with titles show the names of the four most important unit operations. Each box shows the chemical reaction with reaction enthalpy at STP in kilojoules per mole of carbon and the reaction number for reference elsewhere in the paper. Note that water is liberated in reaction 1 and consumed in reaction 4, balancing the process. The full process has evaporative losses, as shown in Figure 2 Our process comprises two connected chemical loops ( Figure 1 ). The first loop captures COfrom the atmosphere using an aqueous solution with ionic concentrations of roughly 1.0 M OH, 0.5 M CO, and 2.0 M K. In the second loop, COis precipitated by reaction with Cato form CaCOwhile the Cais replenished by dissolution of Ca(OH). The CaCOis calcined to liberate COproducing CaO, which is hydrated or “slaked” to produce Ca(OH)

2 /year from the atmosphere and delivers a 1.46 Mt-CO 2 /year stream of dry CO 2 at 15 MPa. The additional 0.48 Mt-CO 2 /year is produced by on-site combustion of natural gas to meet all plant thermal and electrical requirements. Alternate configurations with electricity and gas input are described in the section on Heat and Mass Balance and Alternative Configurations and life cycle carbon balance in the section on Avoided Emissions and Life Cycle Accounting. Figure 2 Overview of Process Showing Mass and Energy Balances Show full caption Electricity demands are indicated in red as MW. Selected gas and liquid streams show the most important constituents using mass fraction as for gaseous streams and molar concentration for aqueous. Mixed phase streams with substantial solid-phase mass flow are color-coded based on the phase of the gas or liquid transporting the solid. Units are indicated with graphical representations that suggest a schematic physical design of the unit. Many minor streams, such as cooling water to the multistage CO 2 compressor, are not shown. As described in the text, there are several options for introducing the fines stream back into the calciner, these are omitted for simplicity, and this heat and mass balance reflects fines being treated identically to the pellet stream leaving the washer. CE has developed a process to implement this cycle at industrial scale. Figure 2 provides a simplified energy and material balance of the complete process (and Figure S1 shows a rendering of one possible configuration of plant equipment to perform this process). At full capacity, this plant captures 0.98 Mt-CO/year from the atmosphere and delivers a 1.46 Mt-CO/year stream of dry COat 15 MPa. The additional 0.48 Mt-CO/year is produced by on-site combustion of natural gas to meet all plant thermal and electrical requirements. Alternate configurations with electricity and gas input are described in the section on Heat and Mass Balance and Alternative Configurations and life cycle carbon balance in the section on Avoided Emissions and Life Cycle Accounting.

Energy and material balances come from an Aspen Plus simulation. That simulation depends on performance models of individual unit operations; and these models depend, in turn, on a combination of vendor data and data from the pilot plant described later.

24 Rickover, H. (1953). Letter. Available at: http://ecolo.org/documents/documents_in_english/Rickover.pdf. Accessed 10 January, 2018. 2 Lackner K.S.

Ziock H.-J.

Grimes P. Carbon Dioxide Extraction from Air: Is it an Option? Technical Report LA-UR-99–583. As with any industrial technology, there is a sharp distinction between the ease of developing “paper” designs and the difficulty of developing an operating plant. To paraphrase Rickover: an academic plant is simple, cheap, and uses off-the-shelf components; whereas, a practical plant is complicated, expensive, and “is requiring an immense amount of development on apparently trivial items.”CE has now spent roughly 100 person-years on such apparently trivial items to develop a process proposed almost two decades ago by Klaus Lackner and collaborators.

25 Baciocchi R.

Storti G.

Mazzotti M. Process design and energy requirements for the capture of carbon dioxide from air. 26 Heidel, K.R., Ritchie, J.A., Kainth, A.P.S., and Keith, D.W. (2014). United States Patent: 8728428-Recovering a caustic solution via calcium carbonate crystal aggregates. Filed March 13, 2013, and issued May 20, 2014. , 27 Heidel, K.R., Keith, D.W., Ritchie, J.A., Vollendorf, N., and Fessler, E. (2017). United States Patent: 9637393-Recovering a caustic solution via calcium carbonate crystal aggregates. Filed May 19, 2014, and issued May 2, 2017. For each unit, we have either identified a vendor of commercial hardware that meets the process specifications or identified commercial hardware that can be adapted to perform the process. In the latter case, we have typically entered into a formal collaboration with a vendor and then tested the unit at a scale the vendor deems necessary to allow specification of commercial-scale hardware. For the major unit operations, this process has involved several cycles of testing at progressively larger scales working with equipment vendors to de-risk the technology. Consider the pellet reactor as an illustrative example of our development process. The idea originated from a paper that suggested use of a Crystalactordeveloped for wastewater treatment by Royal HaskoningDHV (RHDHV). Working with Procorp Enterprises, RHDHV's American licensee, CE developed a different process configuration.The first tests with CE's process conditions were performed in 2011 using Procorp's existing 5-cm-diameter lab unit. CE then contracted with Procorp to build and operate a larger, 30-cm-diameter custom-built system with more appropriate lime injection technology at Procorp's facility in Waukesha, WI. CE then worked with RHDHV and Procorp to design a 1.2-m-diameter system with up to 11 m of fluidized bed depth as part of CE's Squamish pilot plant. Finally, CE built an additional, smaller 14-cm-diameter system at the pilot plant, to speed up testing of alternative operating conditions that are then implemented on the main pilot pellet reactor.

In this section, we first describe the four major unit operations: the contactor, pellet reactor, calciner, and slaker, corresponding to the four reactions depicted in Figure 1 . Performance estimates are based on a combination of data from vendors and from our pilot plant (discussed later), along with data from the minor unit operations (see below). These unit performance estimates then drive a chemical process simulation (see below) that provides the values reported in Figure 2

Contactor 2 from the air occurs at the surface of a ∼50 μm film of solution flowing downward through structured plastic packing through which the air flows horizontally (cross-flow configuration). The transport of CO 2 into the fluid is limited by a reaction-diffusion process occurring in the liquid film with a characteristic e-folding length of ∼0.3 μm. The mass transfer coefficient for CO 2 (K L ) is most strongly determined by [OH−] and temperature. We use a semi-empirical formula to estimate the mass transfer coefficient on representative well-wetted structured packings (the “effective” K L ) for a range of fluid compositions and ambient temperatures, 28 Holmes G. A Carbon Dioxide Absorption Performance Evaluation for Capture from Ambient Air. 15 Stolaroff J.K.

Keith D.W.

Lowry G.V. Carbon dioxide capture from atmospheric air using sodium hydroxide spray. , 25 Baciocchi R.

Storti G.

Mazzotti M. Process design and energy requirements for the capture of carbon dioxide from air. , 29 Tepe J.B.

Dodge B.F. Absorption of carbon dioxide by sodium hydroxide solutions in a packed column. , 30 Spector N.A.

Dodge B.F. Removal of carbon dioxide from atmospheric air. , 31 Danckwerts P.V.

Kennedy A.M.

Roberts D. Kinetics of CO 2 absorption in alkaline solutions—II: absorption in a packed column and tests of surface-renewal models. , 32 Astarita G. Regimes of mass transfer with chemical reaction. , 33 Zeman F. Energy and material balance of CO 2 capture from ambient air. L is 1.3 mm/s at 20°C and a typical solution composition of 1.0 M OH−, 0.5 M CO 3 2−, and 2.0 M K+. The contactor brings ambient air in contact with the alkali capture solution. Capture of COfrom the air occurs at the surface of a ∼50 μm film of solution flowing downward through structured plastic packing through which the air flows horizontally (cross-flow configuration). The transport of COinto the fluid is limited by a reaction-diffusion process occurring in the liquid film with a characteristic e-folding length of ∼0.3 μm. The mass transfer coefficient for CO(K) is most strongly determined by [OH] and temperature. We use a semi-empirical formula to estimate the mass transfer coefficient on representative well-wetted structured packings (the “effective” K) for a range of fluid compositions and ambient temperatures,which integrates our own empirical data and modeling and aligns with previous literature values.A typical Kis 1.3 mm/s at 20°C and a typical solution composition of 1.0 M OH, 0.5 M CO, and 2.0 M K CE's contactor is based on commercial cooling-tower technology, and the design has benefited from close collaboration with SPX Cooling Technologies (SPX), a leading vendor. While the geometry and fluid chemistry differ from conventional cooling towers, CE's design relies on many of the same components, including fans, structured packings, demisters, fluid distribution systems, and fiber-reinforced plastic structural components. 4 Socolow R.

Desmond M.

Aines R.

Blackstock J.

Bolland O.

Kaarsberg T.

Lewis N.

Mazzotti M.

Pfeffer A.

Sawyer K.

et al. Direct Air Capture of CO 2 with Chemicals: A Technology Assessment for the APS Panel on Public Affairs. , 14 Mazzotti M.

Baciocchi R.

Desmond M.J.

Socolow R.H. Direct air capture of CO 2 with chemicals: optimization of a two-loop hydroxide carbonate system using a countercurrent air-liquid contactor. , 25 Baciocchi R.

Storti G.

Mazzotti M. Process design and energy requirements for the capture of carbon dioxide from air. 23 Holmes G.

Keith D. An air-liquid contactor for large-scale capture of CO 2 from air. 34 Keith, D., Mahmoudkhani, M., Biglioli, A., Hart, B., Heidel, K., and Foniok, M. (2015). United States Patent: 9095813. Carbon dioxide capture method and facility. Filed August 21, 2009, and issued August 4, 2015. 23 Holmes G.

Keith D. An air-liquid contactor for large-scale capture of CO 2 from air. The contactor is the heart of CE's air capture technology. It is the unit that diverges farthest from industrial precedent in that cross-flow cooling-tower components are used for a chemical gas-exchange process, rather than the counterflow vertically oriented tower philosophy typically used for chemical processes. This design choice is a crucial enabler of cost-effective DAC, as designs using vertical packed towers are far more expensive.In this paper, we provide only a short overview because the design is described elsewhere.Major differences between our design and common cooling-tower practice include packing depths of ∼7 m rather than the ∼2–3 m common in cooling towers with structured packing, and use of cyclic-pulsing solution flow to minimize pumping energy while maintaining good packing wetting.The air velocity and packing depth are chosen to minimize combined capital and energy cost,and the resulting design parameters summarized in Table 1 Table 1 Summary Data on Major Unit Operations Parameter Value Justification Contactor Process parameters Mass transport coefficient 1.3 mm/s pilot data and laboratory work 28 Holmes G. A Carbon Dioxide Absorption Performance Evaluation for Capture from Ambient Air. Air velocity 1.4 m/s economic optimization of capital and operating costs 23 Holmes G.

Keith D. An air-liquid contactor for large-scale capture of CO 2 from air. Packing specific surface 210 m2/m−3 packing parameters are based on Brentwood XF12560 with pressure drop reduced by 30% (see section on the Contactor) Packing pressure drop 9.7 Pa/m at 1.4 m/s Packing air travel depth 7 m economic optimization of capital and operating costs 23 Holmes G.

Keith D. An air-liquid contactor for large-scale capture of CO 2 from air. Max liquid flow 4.1 L/m2s required for full wetting—manufacturer’s specification Average liquid flow 0.6 L/m2s pilot data on flow rate cycling Performance metrics Fan energy 61 kWh/t-CO 2 ΔP from pilot data and 70% fan efficiency from SPX Fluid pumping energy 21 kWh/t-CO 2 pump efficiency 82% from GPSA data book Fraction of CO 2 captured 74.5% performance model validated by pilot data Capture rate unit inlet area 22 t-CO 2 m−2/year determined from velocity and fraction captured assuming 400 ppm CO 2 Pellet Reactor Process parameters Fluidization velocity 1.65 cm/s pilot and benchtop show good performance at 1.65 cm/s for our target pellet size, performance degrades for significantly lower velocities Bed height 4.5 m rough optimization of cost of managing fines versus cost of increasing retention; optimization uses empirical performance model driven by pilot data Calcium loading 20 kg-Ca/m2hr Pellet size >0.85 mm pellets removed from bed by passing over a 20 mesh (0.85 mm opening) shaking screen Performance metrics Calcium retention 90% performance model driven by pilot data Fluid pumping energy 27 kWh/t-CO 2 determined from loading rate, fluidization velocity, and pumping efficiency of 82% based on GPSA data Calciner Process parameters Bed bulk density 710 kg/m3 pilot data Fluidization velocity 0.25–2.5 m/s minimum and operating fluidization velocity from pilot data Operating temperature 900°C reaction thermodynamics and pilot data Performance metrics CaCO 3 → CaO conversion efficiency 98% pilot data Energy consumption 4.05 GJ/t-CO 2 determined by Aspen Plus simulation in consultation with Technip Slaker Process parameters Pellet water carryover 11% by mass pilot data Operating temperature 300°C estimate based on preliminary tests Performance metrics Power produced from slaking heat 77 kWh/t-CO 2 estimate from simulation, note that the slaker also consumes 32 kWh/t-CO 2 Conversion to CaO 85% estimate based on tests conducted by Ben Anthony at CanmetENERGY Auxiliary Equipment Specifications ASU power usage 238 kWh/t-O 2 quote from major ASU vendor for 95% purity delivered at 120 kPa CO 2 absorber—capture frac 90% Aspen simulation CO 2 absorber—pressure drop 1 kPa Compressor power usage 132 kWh/t-CO 2 Aspen simulation, with validation from independent calculations For each major unit we provide some important process parameters internal to the unit as well as the most important unit performance parameters. Energy consumption values are given for each ton of CO 2 processed by the unit where for calciner, slaker, and compressor, the amount processed is larger than the amount captured from air because of the CO 2 from the power cycle. 2 uptake differs from the evaporation and sensible heat exchange in a cooling tower, as does our use of pulsed flow to maintain a largely stagnant surface film. Indeed, changes in the tradeoffs between fan energy and capital cost alter the optimal design. 23 Holmes G.

Keith D. An air-liquid contactor for large-scale capture of CO 2 from air. Working with packing manufacturers and using computational fluid dynamics simulations performed by Professor John Grace's group at the University of British Columbia (UBC), we find that minor changes to packing geometry can significantly reduce pressure drop while retaining similar mass transfer performance. Pressure drop can be reduced by >30% compared with the Brentwood XF12560 packing we used in the pilot for the same air velocity and surface area density. Improvements on established designs are possible because we are optimizing for different conditions: COuptake differs from the evaporation and sensible heat exchange in a cooling tower, as does our use of pulsed flow to maintain a largely stagnant surface film. Indeed, changes in the tradeoffs between fan energy and capital cost alter the optimal design.Here, we assume that packing in a commercial plant would have a pressure drop 30% lower than XF12560.

Pellet Reactor Carbonate ion is removed from solution by causticization in the pellet reactor (reaction 2). In this fluidized bed reactor, 0.1–0.9-mm-diameter CaCO 3 pellets are suspended in solution that flows upward at ∼1.1–2.5 cm/s. A slurry of 30% Ca(OH) 2 is injected into the bottom of the reactor vessel (where here and throughout slurry compositions are mass fractions). As Ca2+ reacts with CO 3 2− it drives dissolution of Ca(OH) 2 and precipitation of CaCO 3 , but the fraction of Ca2+ that is precipitated onto pellets depends on maintaining a high surface area of pellets relative to the area of circulating fines while minimizing localized high supersaturations of CaCO 3 that form fines. Small seed pellets are added at the top of the bed, and as pellets grow they sink through the reactor until finished pellets are discharged at the bottom. Roughly 10% of the Ca leaves the vessel as fines that must be captured in a downstream filter. The finished pellets are roughly spherical agglomerations of calcite crystals with negligible porosity. This process is adapted from water treatment technology developed by RHDHV, where it is used to remove multi-valent ions such as CO 3 2−. The process was reengineered to allow formation of CaCO 3 pellets in high ionic strength solutions. Our process differs from water treatment in that (1) the causticization agent is the limiting reagent, (2) it uses 30% lime slurry rather than the ∼2% slurries used in water treatment, and (3) the process parameters are optimized to maximize caustic flux per unit bed area rather than water flux. 35 Faust S.D.

Aly O.M. Chemistry of Water Treatment. 3/s providing a solid basis for cost estimates on our plant, which has a flow rate of 166 m3/s. As described above and in the section on the Pilot Plant, our process was developed iteratively using several generations of prototypes. The industrial design draws on RHDHV's experience in engineering and operating large-scale wastewater treatment plants. The high-concentration lime slurry required abandoning the standard dosing racks and adopting a Spiractor configuration,but unlike a free-standing Spiractor vessel, conical feed sections form an egg-carton-like bottom for a large concrete reactors. Similar systems have been used at the Groote Lucht Wastewater Plant and at Bahrain Tubli Wastewater Plants. The Bahrain plant, for example, has a flow rate of 66 m/s providing a solid basis for cost estimates on our plant, which has a flow rate of 166 m/s. 2 capture process on the Kraft kiln off-gases as the minimum-risk baseline technology for an aqueous alkaline DAC process. 4 Socolow R.

Desmond M.

Aines R.

Blackstock J.

Bolland O.

Kaarsberg T.

Lewis N.

Mazzotti M.

Pfeffer A.

Sawyer K.

et al. Direct Air Capture of CO 2 with Chemicals: A Technology Assessment for the APS Panel on Public Affairs. , 22 Zeman F.S.

Lackner K.S. Capturing carbon dioxide directly from the atmosphere. , 36 Lassiter J.

Misra S. Carbon Engineering. 2 uptake kinetics.) Performance gains come from the ability to make pellets, rather than “lime mud,” which is composed of precipitated 10–50-μm-diameter calcite crystals. The pellets are washed and dried easily, removing the need for vacuum filtration, and resulting in pellets that are drier and have lower alkali carryover than lime mud, which in turn allows use of a CFB rather than a rotary kiln. The lack of vacuum filtration and low water carryover reduces energy consumption in the kiln. Moreover, the CFB has lower capital cost than a rotary kiln and it can be oxy-fired. Our choice of pellet reactor and oxy-fired circulating fluidized bed (CFB) are at the heart of the innovations that reduce the capital and energy cost of the DAC process compared with use of a Kraft caustic recovery loop, which accomplishes the same chemical process in Figure 1 . Early process development work at CE and elsewhere considered using the Kraft process followed by a separate COcapture process on the Kraft kiln off-gases as the minimum-risk baseline technology for an aqueous alkaline DAC process.(Kraft processes use a Na while our process uses a K to improve COuptake kinetics.) Performance gains come from the ability to make pellets, rather than “lime mud,” which is composed of precipitated 10–50-μm-diameter calcite crystals. The pellets are washed and dried easily, removing the need for vacuum filtration, and resulting in pellets that are drier and have lower alkali carryover than lime mud, which in turn allows use of a CFB rather than a rotary kiln. The lack of vacuum filtration and low water carryover reduces energy consumption in the kiln. Moreover, the CFB has lower capital cost than a rotary kiln and it can be oxy-fired.

Calciner Calcination of CaCO 3 to produce CO 2 (reaction 3) is accomplished in an oxygen-fired CFB. Our design has been developed in collaboration with Technip's Dorr-Oliver Fluosolids Systems Division from initial design through laboratory testing, CE's pilot plant, and design of the commercial-scale calciner. Technip has deployed high-temperature fluidized beds at comparable scales, including, for example, two 6.7-m-diameter oxygen-blown CFBs used as gold ore roasters in the Goldstrike mine in Nevada. The calciner, along with preheat cyclones, are large steel vessels lined internally with refractory brick. Fluidizing gas is supplied into the bottom of the calciner through a distribution plate made from an arch of refractory, and natural gas is injected directly into the fluidized bed just above the distribution plate using a series of lances. Our conservative heat integration design reduces technical risk compared with alternate designs that maximize energy efficiency at the expense of higher capital cost and technical risk. Incoming pellets, which arrive from the steam slaker at 300°C, pass through two heat recovery cyclones arranged in counter-current configuration with the outgoing gas stream. In the first preheat cyclone, the incoming solids are preheated to 450°C by cooling the outgoing gases from 650°C to 450°C. In the second preheat stage, the solids are further heated to 650°C by cooling the gases from 900°C to 650°C. Following the cyclones, the outgoing gas stream drives a steam superheater, further cooling the gases to 325°C and producing steam for power generation. The outgoing CaO from the calciner is cooled to 674°C in a single cyclone, which in turn preheats the incoming oxygen to the same temperature before the CaO is sent to the steam slaker. 37 Pröll T.

Rupanovits K.

Kolbitsch P.

Bolhàr-Nordenkampf J.

Hofbauer H. Cold flow model study on a dual circulating fluidized bed (DCFB) system for chemical looping processes. 2 (by volume) to the calciner off-gases. The calciner operates at ambient pressure. Leakage of nitrogen into the system is minimized by using steam-fluidized loop seals at the inlet to the steam slaker and between the steam slaker and the calciner. Experience with similar dual fluidized bed systems with intermediate loop sealssuggest that in a worst-case scenario, the steam slaker atmosphere would contain 0.4% air contributing 0.0013% N(by volume) to the calciner off-gases. As our process is, in some respects, derived from the Kraft pulp process, it is useful to compare this calciner with the rotary kilns used in the Kraft process. A single 6-m-diameter CFB of this design will have a capacity of 2 kt-CaO/day. A typical large Andritz rotary kiln calcining lime mud is 5.5 m diameter × 165 m long and produces 1.6 kt-CaO/day. 2 (that is, CO 2 from calcination) to make up for thermal inefficiencies in heating the feed streams and heat losses to ambient air. This makes the calciner approximately 78% thermally efficient, substantially higher than lime mud calciners, which have thermal efficiencies of roughly 39%, though less efficient than limestone calciners, such as the Cimprogetti TWIN-D shaft kilns, which are 89% efficient on the same basis. Our calciner choice and its efficiency are enabled because pellets are easy to dewater and have appropriate fluidization properties. Process parameters are summarized in Table 1 . The minimum required energy to drive the reaction is 3.18 GJ/t-CaO. Our design requires 4.07 GJ/t-CaO equivalent to 5.25 GJ/t-CO(that is, COfrom calcination) to make up for thermal inefficiencies in heating the feed streams and heat losses to ambient air. This makes the calciner approximately 78% thermally efficient, substantially higher than lime mud calciners, which have thermal efficiencies of roughly 39%, though less efficient than limestone calciners, such as the Cimprogetti TWIN-D shaft kilns, which are 89% efficient on the same basis. Our calciner choice and its efficiency are enabled because pellets are easy to dewater and have appropriate fluidization properties.

Steam Slaker 33 Zeman F. Energy and material balance of CO 2 capture from ambient air. Heat from slaking (reaction 4) is used to dry and preheat the pellets, yielding sufficient steam to sustain the slaking reaction.The thermodynamic advantage of steam slaking over conventional water slaking used in the Kraft process is that the slaking reaction enthalpy is released at higher temperatures. Maximum temperature is 520°C for slaking in 100-kPa steam, whereas we operate at 300°C to achieve fast kinetics. 38 Heidel, K.R. and Rossi, R.A.. United States Patent application: 0170327421. High Temperature Hydrator (A1). Filed May 10, 2017, and issued A1. 3 pellets from washing and hot CaO at 674°C from the calciner's oxygen preheat cyclone. Fluidization velocity is 1 m/s, which transports and slakes quicklime (CaO) particles to form Ca(OH) 2 . Small quicklime particles are elutriated and recirculated by a primary cyclone and loop seal, while the much finer slaked particles mostly bypass the cyclone and are captured in a dust collector. The outgoing stream is 300°C hydrated lime, from which sensible heat is recovered and, along with heat from the slaking reaction, used to dry and warm the pellets. Dry 300°C pellets are then fed into a closed-loop pneumatic conveyor driven by recirculating CO 2 and steam to deliver them to the top of the calciner stack. Designed in partnership with Technip, the slaker is a refractory lined bubbling/turbulent fluid bed that is fluidized by recirculating steam flow.It receives ambient temperature CaCOpellets from washing and hot CaO at 674°C from the calciner's oxygen preheat cyclone. Fluidization velocity is 1 m/s, which transports and slakes quicklime (CaO) particles to form Ca(OH). Small quicklime particles are elutriated and recirculated by a primary cyclone and loop seal, while the much finer slaked particles mostly bypass the cyclone and are captured in a dust collector. The outgoing stream is 300°C hydrated lime, from which sensible heat is recovered and, along with heat from the slaking reaction, used to dry and warm the pellets. Dry 300°C pellets are then fed into a closed-loop pneumatic conveyor driven by recirculating COand steam to deliver them to the top of the calciner stack.

Minor Unit Operations 2 generated from power production, the CO 2 compression and cleanup, and the oxygen plant. Key performance characteristics for each of these units are provided in Beyond the four major units described above, the plant requires many additional unit operations that are “minor” in the sense that they present little or no technical risk. This section summarizes configuration of the power island along with the absorber that captures COgenerated from power production, the COcompression and cleanup, and the oxygen plant. Key performance characteristics for each of these units are provided in Table 1 Power Island The power island consists of a natural gas turbine, followed by a heat recovery steam generator (HRSG). We model a GE LM 2500 DLE with a 2 × 1 HRSG configuration. The resulting steam is combined with steam from the slaker, passed through the superheater (to extract heat from the calciner off-gases), and then used to drive a steam turbine that generates the remainder of the power required by the plant. To reduce complexity, our Aspen simulation approximates this using independent steam cycles for the gas turbine and slaker/superheater. After heat recovery, gas turbine exhaust is sent to the CO 2 absorber. CO 2 Absorber The gas turbine exhaust stream is stripped of CO 2 using a conventional counterflow gas-liquid column, using a portion of the fluid stream returning from the contactor. Based on rough optimization using Aspen, we chose a 12 × 7.5 m (height × diameter) column filled with 95 m2/m3 BERL Ceramic packing that captures ∼90% of inlet CO 2 with a pressure drop of 1.08 kPa at an average operating gas velocity of 0.75 m/s. The absorber outlet is ducted to main air contactor where ∼75% of the remaining CO 2 is captured. CO 2 Compression and Cleanup Compression is accomplished using a standard centrifugal compressor. Performance and power demand were simulated on a four-stage centrifugal compressor based on Dresser-Rand data, with a glycol system for dehydration prior to the final stage, and going from atmospheric to 15 MPa at 45°C. The compressor cost estimate included inter-stage coolers and scrubbers, and cost of equipment was generated by Aspen's Capital Cost Estimator and compared with previous vendor estimates. Oxygen Plant We use a conventional cryogenic air separation unit (ASU). Large ASUs are available in multi-train complexes that produce over 30 kt-O 2 /day. Cryogenic ASUs typically produce oxygen up to 99.8% purity and 10 MPa. Cost was estimated by Solaris (see the section on Process Economics) based on a vendor quote for a 1.5 kt-O 2 /day system. Power demand of 238 kWh/t-O 2 for a 120 kPa delivery pressure was estimated by a major ASU vendor.