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An air−liquid contactor for large-scale capture of
CO 2 from air
Geoffrey Holmes and David W. KeithPhil. Trans. R. Soc. A 2012 370, 4380-4403
doi: 10.1098/rsta.2012.0137

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Phil. Trans. R. Soc. A (2012) 370, 4380–4403
doi:10.1098/rsta.2012.0137

An air–liquid contactor for large-scale captureof CO2 from air
BY GEOFFREY HOLMES1 AND DAVID W. KEITH1,2, *1 Carbon Engineering Ltd, EEEL467, 2500 University Drive NW, Calgary,Alberta, Canada T2N 1N4
2 School of Engineering and Applied Sciences, and Kennedy School ofGovernment, Harvard University, 29 Oxford Street, Cambridge,
MA 02138, USA
We present a conceptually simple method for optimizing the design of a gas–liquidcontactor for capture of carbon dioxide from ambient air, or ‘air capture’. We apply
the method to a slab geometry contactor that uses components, design and fabrication
methods derived from cooling towers. We use mass transfer data appropriate for captureusing a strong NaOH solution, combined with engineering and cost data derived fromengineering studies performed by Carbon Engineering Ltd, and ﬁnd that the total costs
for air contacting alone—no regeneration—can be of the order of $60 per tonne CO2 .
We analyse the reasons why our cost estimate diverges from that of other recent reports
and conclude that the divergence arises from fundamental design choices rather than fromdifferences in costing methodology. Finally, we review the technology risks and conclude
that they can be readily addressed by prototype testing.
Keywords: air capture; carbon dioxide removal; cost optimization

1. Introduction
Air capture (AC) is a technology for capturing carbon dioxide from ambient air
in a closed-loop industrial process. AC is one of an emerging set of technologies
for carbon dioxide removal, which includes geological storage of biotic carbon
and the acceleration of geochemical weathering [1]. Capture of CO2 from air
has been commercialized as a pre-treatment before cryogenic air separation. It is
possible to design industrial-scale AC systems using current commercial processes
in a problem-speciﬁc design, but no commercial-scale AC system has yet beenrealized. Arguments about the wisdom of investing in the development of AC
technology turn on estimates of its cost, and of the potential for cost reductions,both of which are hotly disputed. A recent report by the American PhysicalSociety (APS) concludes that AC would cost roughly $600–800 per tonne CO2
using current technologies and suggests that there is little prospect for substantial*Author for correspondence (david_keith@harvard.edu).
One contribution of 12 to a Discussion Meeting Issue ‘Geoengineering: taking control of ourplanet’s climate?’.

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near-term improvement in these costs [2]. A recent analysis by House et al. [3]concluded that the cost would be about $1000 per tonne CO2 . Our engineering
analysis suggests that overall cost will be much lower. Here, our focus is limited tothe design and costing of a gas–liquid contactor for AC. While we do not analyse
an end-to-end AC system, our results are relevant to the full system analysis
because contactor costs account for over half of the APS cost estimate, and we
demonstrate here that contactor cost can be roughly a factor of four lower, using
an alternative design with the same costing methodology.
In this study, we develop a simpliﬁed method for optimizing the design of aliquid–gas contactor for capturing CO2 from air in strong hydroxide solutions,and provide cost and performance data sufﬁcient to enable ﬁrst-order evaluationof contactor cost. We show that costs for contacting alone can be below $60 per
tonne CO2 . A complete AC system requires both a contactor and a system forregenerating the absorbing solution. We do not provide any new analysis of theregeneration cycle here. Finally, this study examines the differences between ourestimate and that provided by the APS. We ﬁnd that the fourfold discrepancy incost estimates for the contactor arise not from differences in costing methodology,
but from the APS’s choice of a design, which, while low-risk, is ill-suited tothe task.
If it can be developed commercially, AC will occupy a technological niche quitedistinct from traditional point-source carbon capture and storage (CCS). AC will
cost more than CCS when both are operated under the same economic conditions,though it may, in some circumstances, compete with post-combustion capture at
power plants because it enables one to exploit the global heterogeneity in costs ofenergy, labour and CO2 disposal. For example, an AC facility operating on lowcost ‘stranded’ natural gas that is able to provide CO2 for enhanced oil recovery at
a location without other CO2 sources might be competitive with post-combustion
capture in high-cost locations such as Canadian oil sands operations.
The central advantages of AC, should it be realized at cost-effective industrial
scale, are as follows. (i) It allows one to apply industrial economies of scale
to small and mobile emission sources—about 60 per cent of global carbon
emissions—which enables a partial decoupling of carbon capture from theenergy infrastructure. (ii) It provides a route to managing carbon emissions
in the transportation sector—where carbon mitigation costs are generally
higher than in the electric sector—by enabling the generation of carbon-neutralhydrocarbons [4]. We note that, without AC, prospects for deep reductions inemissions from transportation are more limited than elsewhere in the economy.
Electric vehicles show great promise but their ability to make deep reductions
across the entire transportation sector is uncertain, as is the potential forlarge-scale hydrogen and environmentally compatible biofuels. (iii) Finally, while
humanity has many options for reducing emissions, none but carbon removal
and solar geoengineering enable signiﬁcant reduction in the overall climate risk,
because risks grow with the accumulating stock of atmospheric CO2 , whichcannot be reduced without industrial-scale carbon-negative technologies such as
AC. Elimination of emissions stops the increase in the climate risk but cannot
meaningfully reduce it on a time scale shorter than a few centuries. Together,
these advantages may outweigh the intrinsic difﬁculty of capturing carbon fromthe air and mean that AC may be important even if it is not competitive with
CCS at large point sources.
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G. Holmes and D. W. Keith
28E

10
28D
28C
28B

30E
38
32

28A

18

32

14
14
30B

32
30

40
15

30D
46

20
12
22
30A
26

16

Figure 1. Carbon Engineering slab air-contactor design.

It is important to note that most of the work reported here has been performed
within Carbon Engineering Ltd (CE), an independent company (based in Calgary,
Canada) that is engineering industrial-scale AC technology. The design data
presented here are based on the performance of commercially available structured
packing and on previously published estimates of the mass transfer of CO2 into
strong hydroxide solutions. New data from an independent engineering study of
our contactor design are presented here, along with CE’s laboratory data, which
conﬁrm these performance estimates when commercial packings are used with
NaOH and KOH solutions in laboratory-scale tests.

2. Liquid-based packed slab air-contactor
Air contacting for direct AC has commonalities with existing gas separation
technologies, and also with large-scale cooling towers. Like an industrial cooling
unit, AC requires ingestion of very large quantities of ambient air, andmanagement of ingested debris and of droplets, ‘drift’, that might be released tothe environment. Our contactor engineering has focused on large forced-draught
units with cross-ﬂow slab geometry, where cross-ﬂow means that the air ﬂow is
orthogonal to the gravity-driven downward ﬂow of the liquid and slab geometry
means that the overall contactor is thin along the air ﬂow axis compared with
its height and length. A representation of CE’s slab contactor design is shown inﬁgure 1.
Here, we present a simpliﬁed cost-optimization model that describes theengineering constraints and cost trade-offs at the heart of our contactor design.
We provide a rationale for the fundamental design choices mentioned earlier. The
cost model is presented in §2a and the justiﬁcation for the most important costand performance parameters is given in §2b–e. The results of the cost model aredescribed in §2f . We close with a discussion of technology risks and uncertaintiesin §2g.
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As mentioned in the introductory section, the exact composition of theabsorbing ﬂuid is a design choice that is determined by co-optimization of thecontactor and regeneration system. The design presented here assumes thatabsorber ﬂuid is an aqueous solution that absorbs CO2 from ambient air withﬂux across the surface of the liquid ﬁlm of order 1 mg m−2 s−1 (see §2e forrationale), and that, under typical operating conditions, each kilogram of solution
absorbs about 20 g of CO2 before it is returned for regeneration. These properties
are typical of a 1–2 M NaOH solution.
Under these conditions, the theoretical liquid-to-gas volumetric ﬂow ratio
required in the contactor is about 1 : 50 000. This ratio is much smaller—less liquid
per unit gas—than is typically seen in liquid–gas contactors used in chemicalengineering applications. This gives us the ﬂexibility to operate the packing
in a lower liquid ﬂow regime than is common in chemical scrubbers and incooling towers. The advantage of low liquid ﬂow rates is that liquid pumping
can be reduced to a minor factor driving system conﬁguration and cost; while
the disadvantage is that low ﬂows increase the difﬁculty of keeping the packing
evenly wetted. We maintain wetting at low average ﬂow rates by varying the ﬂow,
alternating short-period high ﬂow rates that wet the packing and remove dust anddebris with low ﬂow rates that replenish the wetted surface. The manipulationof ﬂow rates to maintain wetting in these conﬁgurations, described in §2d, is an
important part of the CE contactor design.
Economies of scale in the regeneration system drive us to systems that capture
CO2 at a rate of the order of 0.1–1 megaton per year (Mt yr−1 ). We do notpresent the overall system scale trade-off here; rather, we will simply assume thatthe contactor must capture 1 Mt yr−1 and discuss the design of such a system.
At 400 ppm, air is 0.06 per cent CO2 by weight; so a contactor with a 1 Mt yr−1
capture rate must process at least 46 000 m3 s−1 of air (1300 km3 yr−1 ). The choiceof air velocity is one of the many trade-offs between capital and operating costs.
High air velocities increase energy consumption, which is a primary contributor
to operating costs, but they also minimize the required contactor inlet area, andhence reduce the contribution of capital cost to each unit of CO2 captured. The
results of the cost optimization described below show that optimal air velocities
are in the range 1–3 m s−1 .
Note that while AC literature has often mentioned the use of ambient winds,
our engineering has shown that the capital expenditure of equipment contributes
a large enough fraction of the overall capture cost that forced-draught air ﬂow isessential to avoid ‘down-time’ when winds are too low and to allow continuous
operation of the whole process.
Given an air velocity of 1.5 m s−1 , the required inlet area is 38 000 m2 to capture
1 Mt yr−1 , assuming a CO2 capture fraction of 0.75 (representative of the optimal
capture fractions, as we will show below). Perhaps the most conservative design
would be to use multiple packed towers. The advantage of this system is thatthe use of packed towers for gas–liquid exchange is a standard component of
the worldwide chemical industry. This system has no technical risk because suchpacked towers were widely used in commercial systems that removed CO2 fromair with KOH solutions as a pre-treatment for cryogenic air separation. Packed
towers have several serious disadvantages as a basis for a large-scale AC contactor.
The largest packed towers have inlet areas of about 100 m2 ; so hundreds would
be required to capture 1 Mt yr−1 . An array of these towers lacks a mechanism toPhil. Trans. R. Soc. A (2012)

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4384

G. Holmes and D. W. Keithavoid recycling of the low-CO2 outlet air from one unit into the inlet of another,and yet such recycling must be minimized in a cost-effective system. Moreover,the multiple tower geometry is far more expensive because of the cost versus scale
relationships and the complexity of the required piping connections.
The alternative is to adapt technology used in large-scale cooling towers andwaste treatment facilities, which are designed to efﬁciently bring very large
quantities of ambient air into contact with ﬂuids. As described in §2b, our slab
geometry design is based closely on the design experience in these industries.
There are technical risks associated with adapting cooling tower technology to
the chemical process used in AC; we address these risks and the steps we aretaking to minimize them in §2g.
The engineering study described in §2b arrived at an optimized air-contactor
design that is roughly 20 m tall, 8 m deep and 200 m long. In CE’s fullscale facility design, roughly 10 contacting units would be dispersed around acentral regeneration, compression and processing facility, to cumulatively capture
1 Mt yr−1 .

(a) Air-contactor cost optimization model
The simpliﬁed cost model assumes that the contactor’s capital cost can bedivided into two independent components, the shell and the packing. The packing
cost, CP , is proportional to the packing volume and it includes the installed costof the packing itself along with the specialized supports, ﬂuid distributors anddrift eliminators. The shell includes the structure, fans, ﬂuid pumps and largescale ﬂuid piping (but not the distributors themselves, which are included in CP ).
We assume that the shell cost, CA , is proportional to the frontal swept area ofthe packing. As we describe in §2b, these costs are full owners’ cost, includingsite acquisition, all construction and engineering costs as well as interest during
construction.
The core simpliﬁcation of this cost model is that, for a slab geometry contactor,
the shell cost is roughly independent of the depth of the packing. In part, thisis because maximum wind loads play a larger role in determining slab structuraldesign than does the weight of the packing. For example, the cost of a contactor
shell with a 20 × 200 m frontal area does not change signiﬁcantly as the thicknessof the packing is varied from 3 to 15 m.
As discussed in §2e, the CO2 mass ﬂux into the liquid is always proportional
to the concentration of CO2 in the overlying air so that it can be expressed as aliquid-phase mass transfer coefﬁcient KL (with units of velocity) times the CO2mass density in the air [5].
The complete model equations are shown in table 1, with variables deﬁned intable 2. Here, we provide a few notes of explanation and interpretation.
The CO2 captured per unit inlet area of the slab contactor is expressed as amass ﬂux (equation (2.1)). The airborne CO2 concentration decays exponentiallythrough the packing depth with a length scale of V /(3 SSAKL ), which arises
from the ratio of the gas velocity to the CO2 removal ﬂux, which is the masstransfer coefﬁcient times the surface area per unit volume. The packing efﬁciency,
3, accounts for imperfect wetting of the packing area. Note that equation (2.1) isa simple and robust approximation; in reality, hydroxide concentration and hence
uptake rate vary within the packing. We have constructed a full two-dimensional
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A contactor for capture of CO2 from air
Table 1. Cost optimization model equations.
equation

nameformula

units

(2.1)
(2.2)

CO2 capture ﬂux
pressure drop in
packinga
energy for fanscapital cost
operating cost
total cost
minimization

F = fop rCO2 V (1 − e−3 SSA D KL /V )DP = D7.4V 2.14

kg m−2 yr−1
Pa

E = fop DPV /hfan
Ccapital = CA + Cpack D
Coperating = ECelec + M&O Ccapital
min CCO2 = (Coperating + CCF Ccapital )/F

J m−2 yr−1
$ m−2
$ m−2 yr−1
$ per tonne CO2

(2.3)
(2.4)
(2.5)
(2.6)

V ,D

a Vendor-speciﬁed pressure drop for Brentwood Industries XF12560 packing, simpliﬁed to low liquid

ﬂow scenario.
Table 2. Cost optimization variables, values and units.
variable

units

value

notes and typical values

hfan
Celec
CA
Cpack
fop
CCF
M&O
3
rCO2
SSA
KL
V
D

%
$ J−1
$ m−2
$ m−3
s per year
% per year
% per year%
kg m−3m2 m−3
m s−1
m s−1
m

56
2.2 × 10−83700
250
2.7 × 10715
580
7.3 × 10−4210
1.5 × 10−3
1–2
5–15

fan efﬁciency (§2c)
cost of electricity ($80 per MW h)capital cost per frontal area (§2b)
packing and ﬂuid distributor cost (§2d)
assumed annual operation fraction (85%)
capital charge factor
maintenance and operation
packing efﬁciency (§2d)
mass density of CO2 in air at 400 ppmv
speciﬁc packing area (§2d)
mass transfer coefﬁcient (§2e)
air velocity (optimal values computed in §2f )packing depth (optimal values computed in §2f )model as an Excel module in Aspen Plus that accounts for gas ﬂow and depletionof the hydroxide in the liquid in contact with intermittent ﬂuid ﬂow, and we havetested the model in various experimental geometries.
The energy cost of operating the contactor is simply determined fromthe pressure drop (equation (2.2)) and the pressure–volume work term
(equation (2.3)).
To simplify life-cycle analysis, all of our full system designs supply all electricity
needs with on-site generation that is thermally integrated with the regeneration
cycle. Almost all CO2 from the process is captured as part of the regeneration
cycle. In this simpliﬁed contactor-only model, we simply assume a cost ofelectricity of $80 per MW h. In the full design, the two systems must, of course,
be optimized jointly.
We adopt 15 per cent as an overall annual capital charge factor, which includes
interest, depreciation and taxes. The annual cost of operating personnel andregular maintenance is estimated annually as 5 per cent of the system capital cost.
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G. Holmes and D. W. Keith
Table 3. CE 1 Mt yr−1 contactor array, structural cost analysis breakdown.

item

laboura (%)

materials (%)

fans
packing
pumps
structureadmin/contingency
total

3.8
1.4
2.6
27.7

14.8
17.4
1.0
8.5

35.5

41.8

percentage
of costcapital cost
(C$ million)

18.7
18.9
3.5
36.2
22.7
100

26.3
26.6
4.9
50.9
31.9
140.6

a Labour costs include all project management costs as these costs were allocated in proportion tolabour costs.

Finally, the loss of water through evaporation results in the need for a make-up
water stream priced at an average [6] of $0.5 per tonne H2 O, and thus adds ofthe order of $5 to the cost of every tonne of captured CO2 . For simplicity in therest of our analysis, we ignore this cost.

(b) Contactor capital cost
Our estimate of the cost per unit contactor inlet area, CA , derives from a design
study performed by CE and a sub-contractor, on a purpose-built air-capture
contactor, and from cost estimates for industrial cooling tower hardware.
CE contracted the engineering, procurement and construction ﬁrm SolTech
Projects Inc. (Calgary, Canada) in 2009 to develop an engineering design based
upon CE’s slab air-contactor concept. In this initial study, CE speciﬁed packing
performance, including weight, liquid ﬂow rate requirements and pressure drop,
and SolTech was responsible for design and costing of only the shell around thepacking. SolTech performed a full structural engineering analysis and simulated
100-year event wind and snow loading scenarios and examined overall costsensitivity to the height, width and length of the contactor modules. CE’s current
20 × 200 m module size is a result of this structural cost optimization.
SolTech’s ‘bottom-up’ cost analysis and engineering included vendor quotes
for site preparations, foundation, structural materials and construction, ﬁeld
assembly, air and liquid handling systems, as well as standard project
management and contingency. In addition, SolTech performed an independent‘top-down’ factor and scaling analysis of system cost to assess the accuracy of thecost analysis and to account for items not included in this ﬁrst design study.
The full, installed cost of a single 20 × 200 m (note that these are packing
dimensions, external size is larger) CE slab air-contactor module was C$15 million(Canadian dollars). A top-level cost breakdown is presented in table 3.
Note that cost estimation was based on 2008–2009, Alberta-speciﬁc prices andlabour rates, which were among the highest in the world due to rapid growth in
the Alberta petroleum industry. These Alberta costs are typically 30–50% higherthat US Gulf Coast costs. Note that worldwide construction cost indices peakedin 2008. For example, the IHS-CERA ‘Downstream’ capital cost index increased
85 per cent from 2003 to its 2008 peak, from which it has now declined about
20 per cent.
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Table 4. Structural cost comparisons—cooling towers, CE slab contactor, traditional packed tower.

inlet area (m2 )packing volume (m3 )unit cost ($1000)b
packing material
orientation
inlet area cost($ per m2 inlet)

modular counterﬂow cooling unit

large cross-ﬂow
cooling tower

CE slab design(single module)

packed
tower

30
45
125
Brentwood
CF1900counter-ﬂow
4150

1750
10 000a
4000
Brentwood
XF12560
cross-ﬂow
2300

4000
20 000
15 000Brentwood
XF12560
cross-ﬂow
3700

113
285
1770
Sulzer 250X
counter-ﬂow
15 800ca Standard quotes for large-scale cooling towers commonly include this volume of ‘splash ﬁll’ or alesser volume of structured packing (‘ﬁlm ﬁll’), but can easily be ordered with this full volume ofﬁlm ﬁll.
b The unit cost values represent full ‘overnight’ costs. Note that the CE slab cost does include site
preparation, while the cooling tower costs do not.
c Calculated from the APS report using total contactor equipment cost ($290 million) less packing
cost ($160 million), multiplied by their 4.5 scale factor to get installed cost and divided by their
37 000 m2 inlet area.

CE is continuing with design optimization efforts to further reduce costs,
including the exploration of labour cost savings associated with shop fabricationof individual 5 × 5 m contactor units for subsequent ﬁeld assembly. CE also has ajoint development agreement with a leading cooling tower manufacturer to further
examine engineering, construction and materials choices that could further reduceoverall contactor structure cost.
The CE contactor design is closely related to industrial cooling towers. Table 4
compares the results from the SolTech/CE design study with vendor prices
from cooling tower suppliers obtained in 2010, which provide an independent
check on our cost estimates. The ﬁrst system is from a manufacturer ofsmall, modular counter-ﬂow cooling units, and the second is from a leading
manufacturer of large ﬁeld-erected, cross-ﬂow cooling towers, showing that thesefully mature designs with years of optimization and reﬁnement have managed
to achieve inlet area costs of $2300 m−2 . We also compare against a traditional,
enclosed cylindrical counter-ﬂow packed tower, such as those commonly usedin the chemical processing industry, and here we have calculated the costand characteristics from the packed tower design used in the APS report[2, p. 37].
We claim that the designers of slab air-contactors could emulate the design
characteristics and construction methods of these large-scale cooling towers toeventually achieve similarly low structural costs. The inlet area cost for CE’sslab contactor falls in between the costs for comparable systems that ingest large
quantities of atmospheric air, namely the small-scale specialized cooling units, andlarge-scale ﬁeld-erected cooling towers, increasing our conﬁdence in the resultsfrom SolTech’s design study for CE’s slab contactor.
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The working ﬂuid of CE’s slab contactor is indeed liquid hydroxide, rather thanwater, but this difference was handled with minimal cost penalty by the use ofaffordable, hydroxide-resistant, polyvinyl chloride (PVC) packing products andliquid handling components. In fact, one key beneﬁt of the hydroxide working
ﬂuid, which will be discussed later, is that it inhibits biological growth and theresultant fouling problems that are a constant challenge in water-based cooling
tower systems.
Further, the obvious outlier in this inlet area cost comparison was the
traditional packed tower. Although this packed tower design represents the lowest
technical risk scenario, because they are already used for CO2 stripping at
industrial scale, the fully enclosed structure leads to a prohibitively high costand renders this design impracticable for large-scale processing of ambient air.
The technical risk trade-offs associated with moving away from the packed tower
design towards the more cost-effective, open, cross-ﬂow conﬁguration as embodied
by CE’s slab contactor are further discussed in §2g.

(c) Fan and pumping energy requirements
The fans incorporated in CE’s air-contactor design are commonly used inforced-draught cooling towers. These two applications are very similar in termsof required volumetric air ﬂow rate, air velocity and duration of run-time. CE has
selected a Swifter CTX series fan for the air-contactor, which will operate at an
efﬁciency of approximately 66 per cent while moving air through the air-contactor
at approximately 2 m s−1 . There are in fact a number of commercially available
choices for the fans that would perform well in the CE air-contactor because theexpected pressure drop, desired air velocity and volume of air processed are very
similar to forced-draught cooling towers.
The structural study performed by SolTech included a survey of pumps andpiping to determine the necessary pumps and piping diameters for the aircontactor based on the estimated liquid ﬂows. Their analysis showed that,
owing to the intermittent wetting, the liquid ﬂows were so low that the costof pumping equipment became insigniﬁcant, allowing CE relative freedom inchoosing pumping equipment and vendors.
In a typical contactor operating condition, the energy used for the ﬂuid pumps
is approximately 15 per cent of that required for the air fans. For this simpliﬁed
analysis, we simply decrease the fan efﬁciency to 56 per cent from 66 per cent toaccount for this.

(d) Structured packing mass-transfer media
At the core of CE’s air-contactor design is a structured packing material that
distributes the CO2 -absorbing hydroxide solution over a large surface area, andcreates efﬁcient contact with the air by minimizing frictional air resistance inthe ﬂow channels. The large liquid surface area is essential in order to absorb
a large fraction of CO2 from the air, and the low pressure drop is needed tominimize the fan energy requirements at high air throughputs. As we explain inthis section, the material of packing construction and the ﬂow geometry are both
crucial parameters for consideration in a large-scale air-contactor design.
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A contactor for capture of CO2 from air
Table 5. Structured packing speciﬁcations.
parameter

Brentwood XF12560

Sulzer 250X

SSA (m2 m−3 )weight (kg m−3 )maximum temperature (◦ C)
DP/D (at Vair = 2 m s−1 ) (Pa m−1 )cost ($ m−2 )chemical compatibility
materialorientationcompatible demisters

210
40
60
46
1
19 M NaOH
PVC
cross-ﬂow
Brentwood XF150Max

250
200
900a
75
6
19 M NaOH
316SS
counter-ﬂow
Sulzer Mellachevrona High-temperature tolerance is required for many common unit operations.

Traditional packed towers in chemical processing industries commonly use
structured packing products made from stainless steel, which often cost of theorder of $5 m−2 . Note that, while packing products are priced per volume, we
ﬁnd it convenient to express cost per unit surface area to allow comparison
between products with different speciﬁc surface areas (SSAs; see table 5 for acomparison between two packing products). For example, at $1500 m−3 , Sulzer
250X costs $6 m−2 . On the criteria of cost-effectiveness, chemical compatibility
with hydroxide, high surface area, long lifetime, debris management and low
pressure drop, we have identiﬁed a more appropriate family of PVC-based
packing products from the cooling tower industry, upon which to base our slab
contactor design. PVC structured packing products have average costs of less than
$1 m−2 , and so the switch from stainless steel to PVC material produces marked
cost savings for a large-scale slab contactor. Moreover, plastic packings can beconditioned to offer the same NaOH wetting as steel packings, and as we describein §3b,c, the performance of plastic tower packing exceeds that of steel towerpacking in pressure drop per unit of surface area, arguably the single mostimportant efﬁciency metric.
Experiments at both laboratory scale and with a multi-metre tower have been
conducted by CE to quantify the CO2 -absorption performance of both PVC
and steel packing products. Principally, in 2008, a 3 m tall packed tower was
operated with steel packing in a counter-ﬂow mode and with liquid hydroxide, tomeasure the CO2 -absorption performance and fan energy requirements. In 2010,
laboratory-scale tests were conducted with a cross-ﬂow contacting apparatus on
several commercial packing products, made from both steel and PVC. During thistime, CE evaluated the liquid ‘hold-up’ (quantity of liquid retained by a wetted
packing) properties, frictional resistance to air ﬂow (measured as pressure drop)
and CO2 -uptake performance of these products, all under a full range of capture
solution concentrations, air velocities, temperatures, relative humidity values andrun-times (CE’s experimental results are summarized in table 6). Once natural
surface conditioning occurs, as it does in cooling towers, CE has shown that PVC
packing products are able to achieve the same CO2 -absorption and pressuredrop performance metrics as their comparable steel counterparts. We note herethat our recent study presented at GHGT-10 has shown equivalent performance
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G. Holmes and D. W. Keith
Table 6. CO2 absorption by strong hydroxide, literature values.

sourceKL 3(mm s−1 )aSpector & Dodge [7]conditionscomments

1.6

inverted from run 85F(1 inch Berl saddles),using
SSA = 250 m2 m−3 ,D = 3 m,V = 0.66 m s−1 ,CF = 0.84

experimental

Astarita [8]1.8bconverted from other units.
NaOH.
Conditions unclear

converted from a ﬂux
value attributed to
Astarita, reported in
Zeman [9]Herzog [10]0.1bconverted from other units.
1 M Ca(OH)2 exposed
stagnant pool

theoretical

Baciocchi et al. [11]1.2

inverted from table A1,
2 M NaOH scenario,
using
SSA = 250 m2 m−3 ,D = 10.3 m,
V = 4.3 m s−1 ,CF = 0.5

important because thisstudy formed thebasis of the APS
reportStolaroff et al. [12]2.5b2.5 M NaOH

theoretical

Mahmoudkhani
et al. [13]2.3

inverted using
SSA = 250 m2 m−3 ,D = 2.6 m,
V = 0.8 m s−1 ,CF = 0.85

experimental

Zeman [14]1.2

converted from other units.
Random packing,
SSA = 300 m2 m−3 .
1 M NaOH,
V = unreported.
3 accounted forby author

experimental

APS DAC report [2]2.0

inverted from the APS
DAC tower design (APSDAC report [2], table
2.3, p. 35), using
SSA = 250 m2 m−3 ,D = 2.8 m,
V = 2.0 m s−1 ,CF = 0.5

theoretical. Notediscrepancy with
Baciocchi et al. [11](Continued.)Phil. Trans. R. Soc. A (2012)

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A contactor for capture of CO2 from air
Table 6. (Continued.)sourceKL 3(mm s−1 )aCEc

conditionscomments

2.1

inverted from empiricallaboratory data.
XF12560 packing with
SSA = 210 m2 m−3 ,cross-ﬂow, 2 M NaOH,
D = 0.6 m,
V = 1.0 m s−1 ,CF = 0.23

CE laboratory-scale
contactor

2.4

inverted from empiricallaboratory data. Sulzer
500Y packing with
SSA = 500 m2 m−3 ,cross-ﬂow, 2 M KOH,
D = 0.6 m,
V = 1 m s−1 ,CF = 0.49

CE laboratory-scale
contactor

a Unless

indicated otherwise, we have used reported experimental conditions to invert theexponential decay in equation (2.1) to give KL 3 values for this table. We identify which parameters
are used to invert equation (2.1) in each case.
b Reported as K only, not K 3.
L
L
c CE used an in-house laboratory-scale cross-ﬂow contacting apparatus to test Sulzer 250X
and 500Y, and Brentwood XF12560, XF75, CF1200 and CF1900 packing products. Tests wereconducted with CO2 inlet concentrations ranging from 100 to 700 ppm, relative humidity 20–
100%, temperature 20–33◦ C, with air velocity and solution chemistries as deﬁned above. Only two
representative measurements, for 2 M NaOH and 2 M KOH, are reported here.

in CO2 absorption for two PVC and steel packing products with comparable
ﬂow geometries [15]. The long-term performance of PVC packing products withstrong hydroxide solutions will be evaluated by a prototype contactor being
built by CE for long-duration operation in summer 2011 to quantify key risks
as detailed in §2g.
CE’s experimental evaluation shows that commercial cross-ﬂow packing
products—PVC or steel—do not have optimized geometries for AC. Boththe diffusivity of CO2 in ambient air and the mass transfer properties ofstrong hydroxide differ from the conditions normally present in heat-transfer
cooling towers or packed scrubber towers; so there is room for signiﬁcant gains
in contacting efﬁciency with specialized packing design modiﬁcations. CE is
pursuing this development path with suppliers that regularly produce custom
packing designs, but we constrain our discussion in this study to existing
commercial products.
For the purpose of this study, we base our contactor cost analysis on Brentwood
Industries XF12560 structured packing, which is a commercially available productspeciﬁcally designed for use in large cross-ﬂow cooling tower applications, similar
to that of the CE slab air-contactor. XF12560 is constructed from PVC that isPhil. Trans. R. Soc. A (2012)

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4392

G. Holmes and D. W. Keithcompletely resistant to the strong hydroxide, has an efﬁcient cross-ﬂow geometrythat produces low air-side pressure drop and possesses a similar surface area
per volume as the common stainless-steel tower packing Sulzer 250X. The fully
installed cost of XF12560 PVC packing, which includes labour for installation andassociated peripheral products such as supports, mounts, drift eliminators andsolution re-distributors, is $250 m−3 . For a comparison between the characteristics
of Brentwood XF12560 PVC packing product, and the default stainless-steel
product Sulzer 250X, we refer the reader to CE’s GHGT-10 paper [15]. In our
cost analysis, we have used Brentwood speciﬁcations for pressure drop, given as
equation (2.2), and a ‘wetting efﬁciency’ of 0.8, which represents the fraction of
the total geometrical area on which an effective hydroxide ﬁlm forms.
CE has developed a procedure of supplying liquid ﬂow to the packing media
at alternately high and low rates. Brief periods of high liquid ﬂow rates are used
to fully ﬂush the entire packing surface, and subsequently long periods of low‘trickling ﬂow’ are used to conserve pumping energy while allowing the hydroxide
solution to partly react away while capturing CO2 . Through this method, CE’sslab contactor is able to minimize the pressure drop of the packing material by
avoiding air-channel constriction with high liquid ﬂows, and to minimize pumping
energy requirements. In fact, pumping energy requirements are calculated to be
low enough when compared with fan energy usage that they are ignored in our
cost optimization procedure. A discussion of packing-associated risks, including
surface wettability, droplet loss by drift and physical fouling from airborne debris,is included in §2g.

(e) Carbon dioxide absorption by strong aqueous hydroxide
The feasibility of AC using aqueous absorbers rests on the CO2 ﬂux into
the capture solution. This metric dictates the amount of liquid surface area (inpractice, achieved by liquid sprays or structured packing) required behind each
unit of air intake area to absorb a given fraction of ambient CO2 and thus drives
the overall scale and ﬁnancial cost of an air-contactor. We illustrate in this section
that, once normalized to an overall mass transfer coefﬁcient (we use KL , withunits of m s−1 ), literature spanning from the 1940s to the present time reports aconsistent range of experimental and theoretical values for the CO2 mass transfer
performance achievable by strong hydroxide solutions (table 6). The mass transfer
coefﬁcient is crucial to the economics of AC. If KL was an order of magnitude
lower (0.1 mm s−1 ), AC via the wet scrubbing method would probably be
infeasible, whereas if it were an order of magnitude higher (10 mm s−1 ), an AC
contactor would be relatively inexpensive.
Because the ﬂux of CO2 into a liquid hydroxide solution under conditions
relevant to AC is proportional to the mixing ratio of CO2 in the overlying gas,
the mass transfer coefﬁcient therefore has the units of velocity. We denote this as
KL , so that F = rCO2 × KL , where F is the mass ﬂux into the liquid and rCO2 isthe mass of CO2 per unit volume of air at the surface [5].
This capture of CO2 from bulk air via liquid hydroxide occurs in several steps:(i) air-side transport of CO2 to the surface of the liquid ﬁlm, (ii) equilibration withthe liquid surface, (iii) CO2 diffusion inwards through the ﬁlm, (iv) reaction with−OH−(aq) to remove CO2 , and (v) replenishment via diffusion of OH(aq) from withinPhil. Trans. R. Soc. A (2012)

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4393

the bulk ﬂuid and inward diffusion of CO2−
3(aq) . For NaOH, the overall reaction is
2NaOH(aq) + CO2(aq) → Na2 CO3(aq) + H2 O(l) .
It has been shown by others [12] that a combination of the CO2 diffusion andreaction steps (iii and iv) form the limitation to liquid-phase mass transfer, andthat a reaction–diffusion equation [8,16] can be used to model this process:1
vC= DV2 C − C .
(2.7)
vt
t
Here, C denotes the concentration of CO2 in the liquid ﬁlm (in g m−3 ), D is the
diffusivity coefﬁcient for CO2 in the hydroxide solution (in m2 s−1 ) and t is thecharacteristic time constant (in seconds) for the rate-limiting reaction−
CO2(aq) + OH−(aq) → HCO3(aq) .

To calculate our KL , we assume that the ﬁlm of NaOH is stagnant, that theconcentration proﬁle of CO2 in solution is a function of depth only and that theconcentration proﬁle has reached steady state. With these assumptions, we ﬁnd
1
d2 C
− C.
(2.8)
2
dx
t
The simpliﬁed reaction–diffusion equation is solved by using as a boundary
condition the assumption that the liquid-side CO2 concentration at the interface
is kH C0 , where kH is Henry’s constant for CO2 in a hydroxide solution, and where
C0 is the air-side CO2 concentration. The solution yields a characteristic decay
length for CO2 in hydroxide solution (equation (2.9)) and the absorption ﬂux. The
ﬂux expression is normalized with respect to air-side CO2 concentration, to yield
the mass transfer coefﬁcient given in equation (2.10). In these equations, k is thereaction rate constant for the limiting reaction described earlier (in m3 mol−1 s−1 ),and {OH− } is the activity-corrected concentration of hydroxide [12]. Thus we have

D
L=
(2.9)
k{OH− }
0=D

F
= kH Dk{OH− }.
(2.10)
C0
These equations can be evaluated, for a 2 M NaOH solution as an example, using
standard literature values for the parameters. Using values from Danckwerts
et al. [17] for diffusivity (1.21 × 10−9 m2 s−1 ) and Henry’s constant (0.64)
for CO2 in hydroxide, a reaction rate from Stolaroff et al. [12] for thelimiting reaction (8500 l mol−1 s−1 ) and a hydroxide activity at 2 M [OH− ] (0.66)
from Marcos-Arroyo et al. [18], we obtain a characteristic decay length of
0.3 mm, and a mass transfer coefﬁcient of 2.4 mm s−1 . At an atmospheric CO2
concentration of 400 ppmv (0.017 mol m−3 ), the ﬂux into the liquid would be
1.8 mg-CO2 m−2 s−1 .
Our work to characterize the CO2 -absorption performance of strong hydroxide
solutions builds on theoretical and empirical work stretching back to the 1940s.
Table 6 shows the conditions under which previous studies have calculated or

and

KL =

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4394

G. Holmes and D. W. Keithmeasured CO2 absorption. We report KL 3, which is the mass transfer coefﬁcient
multiplied by the packing wetting efﬁciency. The product of these coefﬁcientsdetermines empirically measured CO2 absorption. We focus on the productbecause it is difﬁcult to measure KL and 3 independently in a large-scale sample ofpacking. We obtained the KL 3 values reported here by inverting the exponential
decay term from equation (2.1), in the form 1 − CF = e−SSAD KL 3/V , where CF
denotes capture fraction and is deﬁned by CF = 1 − [CO2 ]out /[CO2 ]in . In our
cost optimization, we use a wetting efﬁciency of 3 = 0.8, but other experiments,
conﬁgurations or packing choices can produce values lower than this.
We note that all but one of the previously reported values, spanning a
multitude of experimental conditions and theoretical parameters over many
decades, are within a factor of three of one another. Given the differences in
research methods used, choices in packing and apparatus, it is remarkable thatthese values are so closely aligned. Herzog [10] is a notable and important outlier
in his choice of liquid solution, and thus mass transfer coefﬁcient, used to evaluate
the feasibility of AC. Herzog used a Ca(OH)2 solution, resulting in a masstransfer coefﬁcient that is an order of magnitude lower than those reported bystudies using Na- or K-based systems with packed absorption apparatus. This
discrepancy is relevant because this study and its author have been widely citedin critiques of the feasibility of AC.
For strong hydroxide solutions, in the 1–4 M range that is most applicableto AC, all of the coefﬁcients in equation (2.10) vary strongly with concentrationand speciation. While the variations in Henry’s constant, diffusivity, reaction rate
and activity with hydroxide concentration can be estimated from the literature,
we ﬁnd that the resulting predictions of KL using equation (2.10) are typicallyinaccurate by factors of order unity. For this reason, we developed a set ofexperiments that produced an extensive database of measured mass transfer
coefﬁcients, as a function of various solution concentrations and temperatures.
Our experiments were carried out under an extensive range of humidity values,
inlet CO2 concentrations, air velocities, packing types and orientation, andsolution ﬂow rates. We use this database in combination with the literature
estimates for the variations of the coefﬁcients in equation (2.10) to produce anempirical function for predicting KL for contactor engineering.
For our baseline cost optimization scenario presented here, we choose aconservative value of KL = 1.5 mm s−1 that is representative of a partially loaded
1–2 M NaOH solution, and apply a wetting efﬁciency of 3 = 0.8, to give KL 3 =1.2 mm s−1 . This is lower than our own measured values by a factor of roughly 50
per cent. It is consistent with the value derived from Baciocchi et al. [11] and is
40 per cent lower than the value used by the APS committee [2] as derived fromthe report in table 6.

(f ) Optimization scenarios and results
Armed with the parameters and values developed in previous sections, we now
discuss the results of our cost optimization model that was itself described in §2a.
We use this simple cost optimization model as a tool to direct our research efforts
towards the factors that offer the most signiﬁcant potential cost reductions. To
illustrate the impact of parameter uncertainty, we choose three simple scenarios:‘base case’ represents our current state of contactor design, ‘pessimistic’ gives
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4395

Table 7. Cost optimization scenarios and results.
parameterabase casepessimistic

optimistic

KL (mm s−1 )CA ($ per m2 inlet)
minimum cost ($ per tonne CO2 )depth at minimum (m)Vair at minimum (m s−1 )

1.5
3700
60
8.6
1.6

0.8
5000
95
14.0
1.5

2.0
2500
43
6.2
1.6

a All other parameters taken at baseline values from §2a.

a conservative upper bound on the cost of atmospheric CO2 contacting, and‘optimistic’ represents positive outcomes of current improvements under way at
CE. Table 7 shows the scenarios and the resulting capture costs.
In each case, the scenario parameters were substituted into equations (2.1)–(2.5),
and the resultant CO2 contacting cost (equation (2.6)) was developed as afunction of contactor depth D, and air velocity V . This function was thenminimized over the domain of depth and velocity to ﬁnd the optimal values ofthese variables that produced the minimal total capture cost. In our base casescenario, the minimized total cost of contacting CO2 from atmospheric air was
calculated as $60 per tonne CO2 , and was produced with a contactor depth of
8.6 m run at an average air velocity of 1.6 m s−1 . Note that the CO2 capture
fraction (approx. 80% in our base case) is not speciﬁed; rather, it is a freedesign parameter value that results from the physical contactor characteristics
that produce the minimum total cost. This is one important difference between
AC, which extracts CO2 from an essentially inﬁnite atmospheric source and seeks
only to do so at minimal cost, and traditional CCS technologies that operate on
ﬁxed gas streams. In this case, the resulting capture fractions are comparable to
CCS facilities, although the design trade-offs that drive the result are different.
Our cost optimization model also calculated the fractional partial derivativesof the minimum cost with respect to the most important parameters. In our basecase scenario, the fractional partial derivative of minimum cost with respect toinlet area cost was 0.54. Thus, with all other parameters held constant, a 10
per cent reduction in inlet area cost would produce a 5.4 per cent reduction in
total capture cost. The fractional partial derivative for mass transfer coefﬁcientwas −0.44; for packing cost, it was 0.27; and for electricity cost, it was 0.19.
Guided by these results, CE is primarily focused on reducing inlet area cost andincreasing mass transfer coefﬁcient to further reduce the total capture cost of ourslab contactor design.
From table 7, it can be seen that our pessimistic scenario produces an upperbound on total contacting cost of $95 per tonne CO2 , and our optimistic scenario
results in a cost of $43 per tonne CO2 . These two scenarios also illustrate theunderlying design trade-offs calculated by our model. In the pessimistic scenario,
with very large capital expense incurred by a $5000 m−2 inlet contactor andthe poor absorptive quality of a 0.8 mm s−1 solution, the optimal operating
regime shifts to a very long contactor depth in order to achieve a high fractional
capture of ingested CO2 , which in turn minimizes the contribution of the capital
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4396

G. Holmes and D. W. Keithexpense to each tonne captured. Alternatively, with a cheap optimistic contactor
running on a highly absorptive 2.0 mm s−1 solution, only 6.2 m of depth are needed
to capture enough of the ingested CO2 to minimize overall cost.
Optimization is crucial to producing a cost-effective contactor design. The
design parameters for AC are sufﬁciently different from the parameters typicalof CCS and of many other industrial gas separation technologies, that adopting
design standards from these technologies without optimization will yield a grossly
sub-optimal air-capture design.

(g) Risks, uncertainties and operational considerations
We consider risks in two broad categories: ﬁrst, environmental health andsafety risks, and second, operational risks and performance uncertainties. Risks tooperators include spills or leaks in piping, but given the century of experiencehandling such solutions in industrial settings, we are conﬁdent that they canbe managed effectively. Risks outside the plant boundary are due to dispersalof ﬁne droplets and the presence of low-CO2 air. These NaOH solutions are
not volatile at ambient conditions; so there are minimal risks from vapourphase transport.
Loss of hydroxide in liquid droplets is the most signiﬁcant environmental risk.
CE’s design aims to reduce loss by two methods. First, and most importantly, thecontactor will include a demister section at each face. Demisters are commerciallyavailable, easily integrated with the packing products we are working with andhave acceptably low pressure drops. For example, a stock demister product
XF80MAx produced by Brentwood Industries assures drift losses below 0.0005 percent of the full ﬂuid recirculation rate, even under the high liquid ﬂow conditions
used in cooling towers as speciﬁcation. Second, our contactors operate in a lowﬂow regime in which there is much less energy for generating small droplets.
During short periods where high-ﬂow operation is needed for debris removal
or ﬁlm establishment, we have the ﬂexibility to reduce air velocity as needed
to control the particle loss rate. We expect that actual overall losses will be
substantially smaller than from standard cooling towers. (Note that this must
be done with care because demisters lose efﬁciency below some air velocity
threshold.)
The central operational risk relates to the long-duration performance of packing
under the low average ﬂow conditions. Risks include fouling with airborne
debris, and, in particular, some interaction of airborne contaminants and packing
surface that might make it difﬁcult to maintain wetting and thus CO2 -absorption
performance.
Most liquid–gas contactors used in the chemical process industry are in theform of closed cylindrical towers operated in a counter-ﬂow conﬁguration withfeed gas interconnections achieved via ducting that couples the gas to otherprocess units. Let us call this a ‘closed’ contactor conﬁguration. These systems arevery different from ‘open’ conﬁgurations used to bring air into contact with anaqueous medium, important examples of which include forced-draught cooling
towers, natural-draught cooling towers, and various sewage and wastewater
treatment applications in which atmospheric air is brought into contact withliquid and micro-organisms growing in ‘bioﬁlms’ are used to treat wastewater.
Traditionally, such open applications used low-technology packing products such
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4397

Table 8. Air-contactor design risks and uncertainties.
risk

closed versus open comparison

fouling

advantage: open. Suppliers of open systems have experience managing
fouling by wind-blown materials; closed systems might requirepre-ﬁlters, and suppliers may lack design experience and mechanisms
for managing fouling.
advantage: closed. Easier to control drift in closed system; however open
systems now have high-performance drift eliminators.
advantage: unclear. Use of strong inorganic base as reagent essentiallyeliminates this risk for either system.
advantage: closed. Designers of chemical separations towers haveextensive experience with a range of corrosives, cooling towers do use
alkalis but not as strong as we require.
advantage: unclear. Larger selection of high-performance packing
available for closed systems, however, cross-ﬂow cooling-towerpackings have performance that exceeds common counter-ﬂow towerpacking (ﬁgure 2)
advantage: unclear. Intermittent liquid ﬂow is needed to minimizing
liquid pumping work. Neither open nor closed contactors nor theirpackings have been designed for this.
advantage: open. Open systems designed to move large volumes ofatmospheric air with low-pressure drops with little re-circulation.
advantage: open. Costs are much lower.

drift losses
evaporative losses
materials compatibility

efﬁciency of packing

intermittent ﬂow

air-side efﬁciency
cost

as ‘splash ﬁll’ media for cooling towers and random (e.g. crushed stone) packing
for trickle ﬁlters. Over the past few decades, there has been a very strong trend
in each of these applications towards use of high-performance plastic-structured
packing products that cost-effectively achieve improved performance.
A useful way to assess the performance risks in the contactor design proposed
here is to consider how the risk might differ for an air-capture facility based on theclosed-system model when compared with one based on technologies derived fromone of the open systems mentioned already. Table 8 provides comparison between
the two across a number of key performance risks. Our view is that—ignoring
cost and considering technology and performance risk alone—there are signiﬁcantadvantages to closed systems, but that there are equally or more important advantages associated with open systems because these systems havepreviously been designed and used to process large quantities of atmospheric air.

3. Comparison with the American Physical Society DAC report
The APS report [2] on direct air capture proposed a reference design for acontactor capturing CO2 out of air using a strong (2 M) NaOH solution withan estimated overall cost for the contactor alone of $240 per tonne CO2 . The
estimated costs presented here are roughly a factor of four smaller. In thissection, we examine the reasons for this discrepancy. We ﬁnd that the cost
discrepancy is driven by design choice rather than by costing methodology. The
APS report used a reference design based on closed counter-ﬂow gas scrubber
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4398

G. Holmes and D. W. Keithcolumn technology—an expensive design choice, and one that makes recirculationof outﬂow gas difﬁcult to manage in a large system with multiple towers. The
APS report suggested that a design such as ours, derived from a cooling tower
design basis, could be signiﬁcantly cheaper, but that design was ruled out of theiranalysis on the basis of technical risk.

(a) Cost analysis methodology
An obvious possibility for the contacting cost discrepancy between our
analysis and that of the APS is that we have made unrealistic economic
assumptions in scaling components to complete systems. Many academic studies
of energy technologies, for example, use unrealistically low capital recovery
factors. However, the APS DAC report and CE use similar costing methods, toscale major equipment costs up to full facility cost estimates. As described in §2b,
our contactor cost estimate used methods consistent with the recommendations
provided in the APS report, appendix A2 [2, p. 43]. In fact, CE uses moreconservative cost assumptions. For example, both CE and APS use a 25 percent contingency, CE assumes a 5 per cent O&M rate, whereas APS assumes
approximately 2 per cent, and, while they are broken down differently, the overall
capital change factor used by APS is 12 per cent while CE’s is 15 per cent. The
APS estimated capital costs by multiplying the raw packing cost by a scaling
factor of 4.5, whereas CE’s costs are based on ﬁrst-order engineering studies
performed by engineering procurement and construction (EPC) ﬁrms undercontract to CE and on vendor quotes from a major cooling tower manufacturer
with which CE has an exploratory joint-development agreement. The results are
consistent in that our bottom-up SolTech study (§2b) did end up with a total
cost that was four times the cost of packing and fans (the major equipment).
Put simply, we concur with the APS’s capital cost estimate for their reference
system, and while APS did not provide a cost for an ‘open’ contactor derived
from cooling tower technology, they do state that such a system ‘would be muchless costly’ (APS [2], section 2.4).

(b) Open versus closed contactor systems
The most important driver of the cost difference between the APS report andthat presented here is that our design is based on an open contactor system,
whereas the APS reference design is derived from closed systems. Both systems
have relevance to AC, with closed systems most commonly used to perform masstransfer on ducted, often toxic, gas streams within chemical processing facilities,
and open systems being the dominant choice for ingesting large quantities ofambient air for cooling applications. As we describe in §2g, there are performanceand development risks associated with both designs. We concur with the APS
that there are risks associated with the use of cross-ﬂow packing media in open
systems with strong hydroxide solutions. We take both the performance and safety
risks seriously, and concur that disciplined design and engineering are required toensure that open air-capture contactors meet existing environmental, health andsafety regulations.
The APS report [2] did not offer a technical analysis of open contacting
systems, but made several references to both their potential and their technicalrisks, namely:
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4399— cyclic liquid distribution to the contactor packing has potential to reduce
pumping costs, but has not been demonstrated (APS, p. 41);
— the practicality of cross-ﬂow designs depends on their ability to provide
good contact between the air and the sorbent (APS, p. 42);— cross-ﬂow designs must minimize losses of absorbent through misting (or‘drift’) or reaction with atmospheric particulates (APS, p. 42); and— cooling tower packing products are less costly than those typically used forgas scrubbing, but have SSAs of 50–150 m2 m−3 and are lower performance(APS, p. 49).
We address these risks in turn, to support our suggestion that open contacting
systems are more relevant for cost and technical assessments of direct AC thanthe closed design chosen by APS.
— Cyclic liquid distribution, whereby a strong ‘pulse’ of liquid ﬂow is used
to fully wet the structured packing and then allowed to slowly react awaywith replenishment from smaller pulses, has the potential to reduce liquidpumping costs in closed or open systems. Preliminary results have beenobtained from cyclic operation in a 2 m tall, 1 m wide, 6 m deep, crossﬂow contactor prototype that we operated outdoors over several months.
Our results have shown that, with pulsing, we can reduce average liquid
ﬂow to roughly 10 per cent of full continuous ﬂow required for adequate
wetting, while retaining an average of 80 per cent of the CO2 capture
rate attained under full ﬂow conditions. Summary data are available
at http://www.carbonengineering.com/wp-content/uploads/2011/12/CE_
OC_Pilot_2011_Results.pdf.
— Good contact between air and sorbent is provided by the choice anddepth of packing material, not by contactor conﬁguration. As long as
they are wetted properly, denser packing products (with higher SSA)
will produce a sharper decay of CO2 concentration through the absorber.
This comes with the trade-off that they also produce a steeper pressure
gradient, and thus the choice of packing density is best accomplishedwithin the overall contactor cost optimization method described in thisstudy. Figure 2 demonstrates that commercial cross-ﬂow packings canperform more than a factor of two better than the packing chosen
by the APS on the ratio of pressure drop to surface area, a crucial
performance metric.
— Drift losses depend on the performance of demisters. Without adequate
demisters, drift losses would be a problem in either a closed or open
conﬁguration. Published data sheets from suppliers demonstrate very
similar demister performance for both cooling tower applications andchemical process towers. The reason is simple; the basic physical design
constraints are precisely the same. In addition, we have demonstrated
with our cross-ﬂow prototype contactor that, through our contactor
conﬁguration and careful installation of commercial demisters, we havecontrolled hydroxide droplet concentration at the contactor outlet to lessthan 10 per cent of the OHSA indoor safety regulations. Again, we refer
the reader to the web-link for a summary of preliminary results.
— We return to the choice of packing product, and its impact on design andcost in §3c.
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G. Holmes and D. W. Keith
1.0
0.9(1/SSA) DP/Dz (Pa)

0.8
0.7
0.6
0.5

Sulzer Y-series (a = 90)
Sulzer X-series (a =60)
Brentwood products
plate flow (dry)

250 Y

0.4

pipe flow (dry)

0.3
0.2

XF12560

0.1
0

200

400
SSA (m2 m–3)

600

800

Figure 2. Pressure gradient divided by speciﬁc surface area (SSA) computed from manufacturers’
data for commercial structured packing products computed at an air velocity of 2 m s−1 and a ﬂuid
ﬂux of 2 l m−2 s−1 (2 l s−1 supplied to each m2 of the top of the packing). Both CO2 absorption
rate and pressure gradient are roughly proportional to SSA; so a crucial ﬁgure of merit for the
performance of structured packings is the pressure gradient divided by SSA. We have plotted thesame ﬁgure of merit for ﬂow in parallel plates and pipes as a guide to the importance of turbulent
drag introduced by real-world structured packings. (Online version in colour.)The preliminary results from CE’s prototype do not resolve the risks mentionedby the APS and discussed earlier, but they do suggest that the technical risks
associated with markedly more cost-effective open contactor designs can bereadily resolved by inexpensive testing.

(c) Packing choice
The APS report [2] states that packing products used for cooling towers
are available with SSAs between 50 and 150 m2 m−3 but that they have lower
performance than those used for chemical process towers; the report does notevaluate their use for direct AC. This choice is important, because the purchasecost of packing was one-third of the total major equipment cost for the entire
APS reference design (APS, pp. 37 and 39); yet plastic packing products similar
to the 250Y model used by the APS report are commercially available at onesixth the cost. Current suppliers of cooling tower packing products do makemedia with SSAs exceeding 200 m2 m−3 , and results presented here and elsewhere
demonstrate that these plastic products can be operated with identical CO2 uptake performance to that achieved by standard scrubber tower packing under
the ambient conditions relevant to direct AC. With all else held equal, a shift fromthe APS stainless-steel packing ($1700 m−3 ) to a cheaper PVC product (approx.
$250 m−3 ) would cut $136 million of bare equipment cost from their reference
design, and would translate to a reduction of $100 per tonne CO2 in their total
capture cost.
In addition to the cost reductions achieved by a switch from steel towerpacking to PVC cooling tower packing, cooling tower designs also offer better
performance in frictional air resistance. We evaluated speciﬁcation sheet data forPhil. Trans. R. Soc. A (2012)

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two lines of Sulzer packing against several Brentwood products, on a basis ofpressure drop per metre, per unit surface area provided, under the air and liquid
ﬂow rate conditions used in the APS report. Our analysis shows that, on a persurface area basis, the Brentwood packing in our design offers less than half thepressure gradient than the Sulzer 250Y used by APS. This translates directly to
lower fan energy requirements, or higher throughput capacity for the contactor,
and thus lower cost.

4. Discussion
We have performed our cost optimization on a slab geometry contactor thatborrows characteristics from gas scrubbing towers commonly used in the chemical
processing industry, and from cooling tower designs that are well suited tocost-effective bulk air processing. Because of the strong trade-offs between capitalcosts and operational energy requirements in air-contactor design, we claim thata ﬁrst-order cost optimization is essential to achieve sensible cost estimates. We
ﬁnd that the cost of air contacting can be of the order of $60 per tonne CO2 witha contactor design that results in a capture fraction of 80 per cent. This capture
fraction stands in sharp contrast to some discussions in AC literature that assume
values above 50 per cent to be cost-ineffective owing to the increasing marginal
energy cost of capture as capture fraction increases. Our optimization, however,illustrates that the additional energy cost associated with this higher CO2 capture
fraction is warranted by the offset of amortized capital expenditure achieved by
the increased capture rate.
Further, when coupled with realistic costs for the regeneration system thatmust complement the air-contactor to form a full direct air-capture facility, evenif we use regeneration costs from the APS report itself, our cost estimates for
AC are much lower than those in that report [2]. On the basis of our rough deconstruction of the APS cost analysis, the total cost for their regeneration system
is $180 per tonne CO2 captured. Coupled with our result for contactor cost of $60
per tonne CO2 captured, and scaled to cost per tonne avoided by the same method
as used by APS, this produces a total cost estimate for a full slab contactor DAC
system of $343 per tonne avoided. (Note that, while we adopt the APS’s costing
methodology here for comparative purposes, we note that the APS’s decision not
to use on-site electricity with integrated capture increases the cost of AC for theirreference system. Standard life cycle analysis methods support the view that one
should count emissions associated with materials that cross the system boundary,but not with electricity generated and consumed inside it [19].)
The APS study stated that [2, p. 42]:
The largest uncertainty in the DAC cost is the cost of the air contactor. It would be muchless costly if an operable process could be designed with a more open system, such as is usedin cooling towers. The practicality of such a system depends on the ability to provide good
contact between the air and the sorbent, while minimizing physical losses of the absorbent
solution through mechanisms such as misting or chemical reactions with particulates andacid gases in the air.

We have suggested that alternative contactor designs can achieve much lower
costs than the packed tower reference design chosen by APS, and thus much
lower overall costs for direct AC. Technical risks of open-air contactor designs
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4402

G. Holmes and D. W. Keithdo, of course, remain to be addressed. But because contactors are inherently
modular, these risks are readily studied by testing at surprisingly small scales. At
Carbon Engineering, we recently ran a 1 m wide by 2 m tall by 6 m deep crossﬂow contactor prototype for several months, for a total cost of order $1 million.
Our preliminary data indicate that we have indeed addressed many of the key
risks raised by the APS report. Given that a single full-scale 5 × 5 m module ofCE’s slab contactor could be built and piloted for under $10 million, resolutionof disagreements about the costs and risks of AC will be resolved once relatively
simple and high-value tests are conducted and independently reviewed.
Finally, in the context of a special issue on geoengineering, we have been asked
by the editor to comment on how AC might ﬁt in to the portfolio of technology
options for managing emissions and climate risk. We see AC as a tool for emissions
mitigation. Carbon Engineering is focused on using AC to enable production oftransportation fuels with low carbon intensity, and, in that context, we see it ascompeting with alternative transportation fuels such as electricity, biofuels andhydrogen [4]. Like any large-scale industrial facility, it poses local risks, along witha global beneﬁt in the form of low-carbon fuels, but these risks are similar in kind
to risks from other energy technology, such as biofuel reﬁneries. We anticipate
that existing health, safety and environmental regulations can and should be
used to regulate the risks of such technologies.
This is in sharp contrast to solar geoengineering technologies, which pose
profound global risks and beneﬁts, requiring new mechanisms of global governance.
One of us (DWK) has argued that the costs of such technologies are sufﬁciently
low (and the risks they pose are sufﬁciently high) that development should bedirected only by public entities acting in the public interest. Development of solar
geoengineering technology by private for-proﬁt companies should be discouragedor, where feasible, banned outright.
The next major step in the development of AC is for companies such as Carbon
Engineering to build and operate several early-stage commercial pilot projects totest the real-world costs and performance of AC. The development of such pilots
under appropriate independent regulatory oversight should go a long way towards
resolving uncertainties about cost and local environmental risk.

5. Summary
We have presented a conceptually simple method for the optimization of aslab geometry air-capture contactor. Design parameters for AC are sufﬁciently
different from the parameters typical of CCS of many other industrial gas
separation technologies that adopting design standards from these technologies
without optimization will yield a grossly sub-optimal air-capture design. With
standard industrial costing methodology applied to our slab contactor structureby an independent EPC ﬁrm, and under very similar economic assumptions as
used by the APS DAC report, our optimization method produces total contactor
costs of the order of $60 per tonne CO2 . Owing to the trade-off between capitaland operating cost, the total cost is not simply linear in either major equipment
cost or energy requirement. Finally, we note that the fourfold discrepancy between
our estimate of contactor cost and that in the recent APS DAC report is due
to fundamentally different design choices, insufﬁcient optimization in the APS
design and our choice of lower-cost contactor internals.
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Both authors are employed by, and have ownership in, Carbon Engineering.
The authors thank the entire Carbon Engineering team for their collaboration and support: Kenton
Heidel, Dr Arvinder Pal Singh, Dr Jianjun Dai, Dr Matt Henderson, Kevin Nold, Hossein Safaei
and Experience Nduagu.

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