CO2 Capture Using Fluorinated Hydrophobic Solvents - Industrial

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CO Capture using Fluorinated Hydrophobic Solvents Paul Mobley, Aravind Rayer, Jak Tanthana, Thomas R. Gohndrone, Mustapha Soukri, Luke J.I. Coleman, and Marty Lail Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b03088 • Publication Date (Web): 22 Sep 2017 Downloaded from http://pubs.acs.org on September 25, 2017

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Industrial & Engineering Chemistry Research

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CO2 Capture using Fluorinated Hydrophobic

2

Solvents

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Paul D. Mobleya, Aravind V. Rayera, Jak Tanthanaa, Thomas R. Gohndronea, Mustapha Soukria,

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Luke J. I. Colemana, and Marty Laila*

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a

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*Corresponding Author: [email protected]

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Keywords: amine, solvent, carbon capture, carbamate, heat of absorption, vapor-liquid

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equilibrium, reboiler heat duty

RTI International, 3040 Cornwallis Road, Durham, North Carolina, 27709

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Abstract: Finding more efficient gas-liquid scrubbing systems with lower parasitic energy

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penalties is important for the future deployment of carbon capture plants for large point source

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CO2 emitters. Minimization of the energy penalty using advanced solvents is one way to reduce

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the energy penalty. Non-aqueous, hydrophobic solvents are one type of solvent in which the

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physical properties of the solvent combined with low heats of absorption and low loading at high

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temperature even with high CO2 pressure result in promising solvents with low estimated

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reboiler heat duty. In this paper, a solvent composed of a hydrophobic amine (2-

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fluorophenethylamine) combined with an acidic, hydrophobic alcohol (octafuoropentanol) is

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studied mechanistically and the experimentally determined reaction products, heats of

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absorption, and vapor liquid equilibria are reported. Approximating process models are

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compared indicating the potential to lower reboiler heat duty in a commercial implementation.

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Introduction

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Finding energy efficient gas-liquid absorption systems for post-combustion CO2 capture is

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essential to achieving commercial acceptance and widespread deployment of Carbon Capture

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and Sequestration (CCS) in the power sector and other industrial sectors with high CO2

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emissions. While technologies already exist and are being tested at large scale which can be used

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to minimize post-combustion CO2 emissions from fossil fuel fired power plants or other large

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industrial point sources such as cement or steel plants, they are currently accompanied by a high

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cost in terms of the energy or power which must be paid to separate CO2.1 This so-called

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parasitic energy load or penalty leads to overall higher fossil fuel consumption in order to obtain

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energy with reduced carbon emissions. Minimization of the parasitic energy penalty using

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advanced solvents will minimize the power deficit and associated cost as well as the additional

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fossil-fuel consumption.

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In a conventional post-combustion carbon capture plant, a solvent is used to absorb CO2 from an

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exhaust gas which contains approximately 5-15 kPa CO2 at about 40°C with thermal swing

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being used to regenerate the solvent and produce a concentrated CO2 stream. The equation (see

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eqn. 1) governing the amount of heat input needed to reverse CO2 binding in the solvent

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includes, at a minimum, sensible heat and heat of absorption.2 If the reactive component in the

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solvent requires a stripping agent to lower the CO2 partial pressure to maintain off-gassing in the

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regenerator, then the heat duty equation also contains the heat of vaporization. The use of


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hydrophobic, non-aqueous, organic components as main solvent constituents can be beneficial

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with regard to sensible heat due to their physical properties. Water has a higher specific heat

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capacity among liquids, while many organics have a specific heat capacity that is considerably

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lower. This can lead to a small reduction in the sensible heat required to heat the solvent to the

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regeneration temperature. However, the advantages mentioned above must be counter-balanced

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with the molecular weight of the non-aqueous solvent components, which are always much

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heavier than water and most amines used in conventional aqueous systems. The molecular

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weight effect mitigates much of the advantage that might seem to benefit solvent components

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with lower specific heat capacities, but typically there is a small improvement. Non-aqueous

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solvent components with lower molecular weights are often not viable as solvent components

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due to their high volatility and presence as volatile organic emissions in the cleaned flue gas.

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= Reboiler Heat

∆

∗ ∗ + ∆, ∗

Sensible Heat

!

∗

∆"#,

+

Heat of Vaporization

(1)

Heat of Reaction

55 56

If a solvent does not require a stripping agent in order to promote degassing in the regenerator

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then there can be savings in terms of the heat of vaporization required for regeneration relative to

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solvents which do require a stripping agent. The heat of vaporization is a variable contributor to

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the regeneration heat of aqueous systems depending on circulation rate, steam supply rate, and

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specific solvent blend, but can contribute roughly 40-60 percent or more of the heat

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requirement.3-4 In order to avoid the use of a stripping agent the absorbing compound must show

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the appropriate CO2 vapor-liquid equilibrium isotherms. Because amines are good

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chemisorbents, many absorb CO2 too strongly to be operated without stripping agents, and this is


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the case for many of the hydrophilic amines. However, a few hydrophobic amines have been

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identified which absorb CO2 readily at low partial pressure (~0.15 bar) and absorber

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temperatures (40°C) while releasing it readily at regenerator temperatures and elevated pressure

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(> 2 bar). The hydrophobic amines of this type have the potential to work either with or without

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a stripping agent and are lower in energy penalty in part because the heat of vaporization

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contribution is small.

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The heat of absorption of the reaction between CO2 and the amine is the largest contributor to the

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heat requirement for non-aqueous solvents. This appears to be the area where hydrophobic

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amines differ somewhat from hydrophilic amines. Typically, the heat of absorption of CO2 is

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measured at several different temperatures, with the heat of absorption at the regeneration

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temperature being the heat that must be input to the process to desorb the CO2 from the solvent.

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For most aqueous solvent systems the heat of absorption is measured to be somewhat higher at

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regeneration temperatures (100-120°C) compared to the heat of absorption at absorber

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temperatures (40°C).5

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In the literature, there are numerous examples of researchers who have investigated non-aqueous

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solvents for CO2 capture. Broadly these can be divided into solvent systems which proceed by

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CO2 capture mechanisms which are carbonate forming (including alkyl, aryl, and bicarbonates)

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and those which operate by amine carbamate forming mechanisms. This paper focuses on non-

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aqueous solvents which operate by the amine carbamate forming mechanism and include non-

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aqueous diluents as components in the solvent systems. Numerous carbamate-forming non-

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aqueous solvents have been investigated in the past, with a notable distinction being the ones

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which are hydrophilic from the ones which are hydrophobic.


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Hydrophilic non-aqueous solvents are more commonly studied than the hydrophobic solvents.

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These hydrophilic systems include the use of alcohols,6-9 glycols,10-13 methyl ethyl ketone,14

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acetonitrile,14 DMSO, glymes,15 or chloroform.16 In order to compare the non-aqueous solvents

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to aqueous amine systems, there are detailed studies on the reaction mechanism,16-17 the reaction

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kinetics,8, 14, 18-19 and CO2 absorption solubility.10-12 Particular attention is paid to the physical

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properties such as decomposition temperature, viscosity, and vapor pressure of the mixture.

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Methyl ethyl ketone, glymes, and acetonitrile are all aprotic polar solvents that will interact less

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with the carbamate product than alcohols, which are polar protic solvents. It is found that the

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choice of solvent can influence the reaction pathway and thus the overall CO2 solubility. A

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solvent such as DMSO interacts with primary amines to promote the formation of carbamic acid,

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while solvents such as glycols promote the formation of the carbamate species.17 The carbamate

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formation follows the 1 mole of CO2 per 2 moles of amine as described extensively in the

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literature. The CO2 solubility is higher when there is carbamic acid formation.20 An

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intermolecular deprotonation occurs after the zwitterion is formed, and then the solvent can

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stabilize the carbamic acid through hydrogen bonding.20 In many cases with non-aqueous

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solvents, the CO2 solubility has been seen to increase above the expected 1:2 reaction

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stoichiometry.

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Deep eutectic solvents are another example of utilizing a polar solvent as a diluent for

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amine/CO2 capture.21 Although the CO2 solubility is relatively high and the solvent interactions

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can promote the carbamate pathway, the viscosity of the solvents and amines are very high (100

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– 300 cP at 25°C. Similarly, Perry et. al.22 designed aminosilicone solvents for CO2 capture

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applications, that utilize amine chemistry to form the carbamate product and suffer from similar

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viscosity limitations.


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The study of hydrophobic amine systems is much less common. One of the apparent drawbacks

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of these systems is that the reaction product, the amine carbamate, is not always soluble in the

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hydrophobic solvent at high concentrations. The solvent and the amine must be carefully

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designed to assure that each are miscible with each other and that the reaction product is

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miscible. If the product is not miscible, it leads to the precipitation of the product and the

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creation of a liquid and solid phase. In order to desorb the reacted CO2 and regenerate the solvent

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in the situations where a phase split occurs, it would be necessary to design a system that can

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economically regenerate the solvent from a two-phase system instead of the well understood

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one-phase system that is observed in most amine/CO2 desorption columns.

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The purpose of using a hydrophobic solvent is to limit the amount of water that is absorbed in the

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process, minimizing the contribution of the heat of vaporization in the regeneration process.

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Primary and secondary amines have been studied in hydrophobic organic solvents such as

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toluene and benzene.14, 16, 18 Kortunov et. al.17 used NMR studies to analyze the reaction of CO2

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with primary and secondary amines in many organic solvents. From their results, they found that

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less polar solvents, for example toluene, would not stabilize the carbamic acid through hydrogen

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bonding and they would promote the formation of the carbamate salt and the 1:2 reaction

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mechanism.

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Room temperature ionic liquids (RTILs) are well studied for the use as CO2 capture solvents that

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physically absorb CO2 and can be designed to be hydrophobic diluents with primary and

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secondary amines. Camper et. al.23 diluted MEA in an imidazolium based IL paired with the

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hydrophobic TF2N anion. They found that the MEA/IL solution had similar capacity to that of


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the aqueous MEA system; however, they experienced issues with the solubility of both primary

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amines in the RTIL and solubility of the carbamate product.

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Besides the physical and thermodynamic properties which are directly related to the regeneration

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heat duty, there are three others which govern the process-design, namely viscosity, volatility,

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and hydrophobicity. The viscosity of all non-aqueous solvents tends to be higher than aqueous

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solvents due to the extremely low viscosity of water. Several classes of non-aqueous solvents

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such as neat ionic liquids and switchable ionic-liquids have such high viscosities that they are not

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feasible for flow in absorber columns and can suffer from mass-transfer limitations. Neutral non-

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aqueous solvents can have viscosities that are only slightly higher than water, which slows

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kinetics, but is still feasible for flow in columns using conventional gas treatment packings.

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The volatility of CO2 solvent components is of great concern and this is of special importance for

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non-aqueous solvents due to both volatile organic emissions concerns as well as concerns about

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the economics of replacement or make-up solvent. As a rule of thumb the volatility goes opposite

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viscosity. The more viscous ionic liquids have negligible vapor pressure and represent the best

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case for avoiding fugitive solvent emissions. The less viscous, neutral, non-aqueous solvent

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components have higher vapor pressures, but if chosen judiciously, can be used without

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significant losses in the treated gas. Almost all volatiles will be recovered from the regenerator

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prior to CO2 compression as a requirement before introduction of CO2 into a pipeline.

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Hydrophobicity, or minimization of the amount of water which can be absorbed as a single phase

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with the non-aqueous solvent, is an impactful property which should be kept low to optimize the

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performance of most non-aqueous solvent systems. The rationale for keeping a low

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concentration of water in the system is two-fold, first to avoid any unwanted and unneeded heat


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of vaporization and second to minimize any hydrogen bonding which might otherwise stabilize

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the CO2-amine complex and lead to higher heats of absorption.

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Insight on how to lower the parasitic energy penalty of a conventional aqueous stripping system

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can be found in the literature.2 There are examples of improvement by structural change in the

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amine as well as through pure process improvement. One important factor is the reboiler heat

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duty, which in a conventional configuration, is responsible for the largest part of the cost (56%).1

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By looking at the reboiler heat duty in a conventional system one can think outside the confines

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of aqueous solvent systems given certain modifications, primarily pertaining to the heat of

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vaporization that is required for a conventional steam stripper. If a non-aqueous solvent can be

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operated without using steam for stripping, or preferably without using any stripping agent

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whatsoever, there is the potential to bring the parasitic energy penalty down from the levels

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presently estimated by the CCS community.

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Hydrophobic amines and hydrophobic organic diluents exist over a range of overall low water

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solubilities. In order to minimize the amount of water solubilized in the solvent, thus minimizing

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potential heat of vaporization and sensible heat contributions during regeneration, we selected a

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fluorinated amine and a highly-fluorinated alcohol as formulation components. In this paper, we

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report on a hydrophobic, non-aqueous solvent system consisting of 2-fluorophenethylamine and

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2,2’, 3,3’, 4,4’, 5,5’-octafuoropentanol (2-FPEA/OFP, Formulation 1), which captures CO2 as a

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conventional amine carbamate with the octafluoropentanol being a spectator or playing a

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hydrogen bonding role with respect to the reaction chemistry while playing a beneficial role to

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the overall solvent performance.

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Experimental


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The system which we report is an equimolar mixture of hydrophobic fluorinated amine in

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hydrophobic

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fluorophenethylamine in 2,2’, 3,3’, 4,4’, 5,5’-octafluoropentanol. The equimolar stoichiometry

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of the components was determined experimentally to avoid precipitation of solid amine

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carbamate after reaction with CO2. Reagent grade 2-fluorophenethylamine was purchased from

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Alfa Aesar and used without further purification. 2,2’, 3,3’, 4,4’, 5,5’-octafluoropentanol was

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purchased from Sigma Aldrich and used as received. Carbon dioxide (bone dry grade, 99.9%)

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was purchased from Airgas. NMR experiments were performed on a Bruker 300 MHz

183

spectrometer. NMR spectra were acquired in CDCl3 solvent. 1H NMR peaks were referenced

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against the residual CHCl3 solvent peak (δ, 7.27 ppm). 19F NMR peaks were not referenced but

185

were evaluated qualitatively to establish reaction pathway.

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Viscosities are measured with a Brookfield DV-II+Pro Viscometer. The Brookfield viscometer is

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equipped with a ULA-304 s/s spindle to measure viscosities in the range of 1 – 2,000 cP and the

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temperature is controlled between 40 – 80°C with a jacketed cell and external circulating bath.

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The viscometer requires ~15 ml of sample to fully immerse the spindle allowing for an accurate

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measurement of the viscosity. The spindle speed (rpm) was set so that the torque measured is 10

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– 100% of the maximum calibrated torque for the spindle. The accuracy of the viscometer was

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checked with the Brookfield Fluid 100 viscosity standard (97.0 cP @ 25°C).

acidic

fluorinated

alcohols.

Specifically,

the

solvents

consist

of

2-

193 194

The vapor-liquid equilibrium (VLE) and heat of absorption experiments in this work were

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performed in a unit composed of a ChemiSens CPA-102 Reaction Calorimeter and an automated

196

gas handling system, including injection and pressure monitoring operating in a batch injection

197

mode. The calorimeter is housed in an incubator held at 30°C to reduce any effects on the


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system arising from changes in ambient conditions. The reaction calorimeter is a highly accurate,

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heat-balanced batch reactor system capable of observing thermal changes on the order of ± 0.001

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J. The reactor vessel was designed to operate at temperatures ranging from 0°C to 200°C and

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pressures ranging from 0.1 kPa to 15 bar with a working sample volume ranging from 10 mL to

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150 mL in a 250 mL reaction vessel. Similarities between the calorimeter reactor vessel and

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equilibrium cell design requirements enabled the use of the calorimeter with an automated gas

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handling and pressure monitoring system to also perform vapor-liquid equilibrium

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measurements. Pressure transducers with full-scale ranges of 0.34, 2.07, and 10.34 barg were

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used to provide highly accurate data at different conditions, each with an accuracy of 0.03% of

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the full-scale output. For measurements above 85°C, a 6.89 barg pressure transducer capable of

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compensating for temperatures up to 135°C was used with an accuracy of 0.05%. The system

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includes two differently sized batch vessels to supply the system with small or large amounts of

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CO2 for VLE measurements (150 mL), and for heat of absorption measurements (1 L).

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In a typical heat of reaction measurement, all process connections were made with an empty,

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clean reactor cell installed and the system was then pressure tested using high-purity N2 to ensure

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that the system was leak proof. Subsequently, the system was repeatedly degassed by vacuum

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(0.138 bar) for a short period and refilled with ultra-high-purity N2 (> 99.999% pure) to evacuate

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the system of any other gas species. Following five cycles of vacuum and N2 addition, the system

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is vented to maintain 1 bar of pure N2 in the cell. Subsequently, a known quantity (mass and

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volume) of approximately 100 mL of amine solvent was loaded through a septum into the

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calorimeter vessel by a cannula from a N2-pressurized round-bottom flask. The vessel agitator

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was started, and the equilibrium cell was allowed to thermally stabilize at the desired absorption

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temperature in a N2 environment. The partial pressure of the nitrogen remaining in the vessel and


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the vapor pressure of the solvent at the absorption temperature were accounted for in the analysis

222

of the data. Typical vapor pressure values for 2FPEA and OFP are 0.11 kPa and 0.79 kPa at

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25ºC, respectively.24-25 The temperature of the solvent was measured and controlled by a highly-

224

accurate and stable Pt-100 RTD, which has an accuracy of ± 0.1°C. CO2 is introduced to the cell

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and the amount of heat generated by the absorption of CO2 is determined by performing a system

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wide, continuous heat balance.

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The quantity of heat generated by the exothermic absorption of CO2 was measured by

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maintaining the thermal stability of the equilibrium cell, by extracting the heat of CO2 absorption

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via an externally controlled and circulated thermal fluid. The system was calibrated with a

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known heat input comparable to the heat released during absorption (i.e., 2 W) for one hour and

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then removed, and allowed to thermally stabilize again. The loading of CO2 into the solvent was

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completed during one continuous, flow-rate-limited injection (maximum of 30 sccm), and

233

discretized during analysis. This method was found to greatly reduce the uncertainty associated

234

with performing the heat of absorption measurements with individual injections. At the start and

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finish of each injection, the cascade PID control of the thermostat unit of the system often results

236

in overshoot, which must be integrated into the measurement. While performing heat of

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absorption measurements with individual injections, it appeared that the operator could skew the

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data to give the appearance of more well-behaved data by changing the bounds of integration and

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the baseline of integration. This potential for data-fitting is removed by performing one slow

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injection. The baseline thermal load is calculated from the average from 30 minutes before the

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start of the injection. The operator selects when the thermal load has stabilized following the

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completion of the injection and adjusts the final baseline thermal load as necessary. A linear

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interpolation between the initial and final baseline thermal load is used and the time between


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broken into approximately 20 “batch injections”. The heat of CO2 absorption is calculated for

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each batch injection experiment using the equation below: ∆$%& =

$%& '$%&

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∆Habs: Heat of CO2 Absorption [kJ/mol CO2]

247

Qabs:

Integrated heat released during CO2 absorption [kJ]

248

nabs:

moles of CO2 absorbed [moles]

249

The heat of absorption measurements were completed differential in temperature and semi-

250

differential in loading. Once the vapor pressure reached the specified maximum, the system was

251

allowed to thermally stabilize and the system was once again calibrated by adding a known heat

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input for one hour and then removed for one hour before shutting down the system. The method

253

was verified using an aqueous blend of thirty weight percent monoethanolamine.

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For VLE experiments, known amounts of ultra-high purity CO2 (> 99.999% pure) were added to

255

the reactor vessel intermittently. The predetermined quantity of CO2 was introduced to the

256

thermostated equilibrium cell in flow-rate-limited injections (maximum of 30 sccm) from a

257

temperature controlled, pressurized vessel. The predetermined quantity of CO2 introduced to the

258

equilibrium cell per injection was based on a CO2 loading of 0.025 mol CO2 / mol absorbate.

259

Once the injection of CO2 was complete, the pressure in the equilibrium cell was allowed to

260

stabilize. The system pressure was considered stable and the system at equilibrium once the

261

deviation in pressure was less than 0.0775 kPa per 30 minutes over a 1.5 hour period. Once the

262

system stabilized, another predetermined quantity of CO2 was introduced to the thermostated


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263

equilibrium cell and the process was repeated until a specified pressure in the reactor was

264

exceeded once equilibrated. The quantity of CO2 absorbed by the solvent for each injection was

265

determined by the following calculations ,-,()(. ,-,1()$2 4'()* = + − 3 1 /-,()(. /-,1()$2 56'9 = +

,9,1()$2 ,9,()(. :49 − 4&;2,2 < − 3 2 569 /9,1()$2 /9,()(.

'$%& = '()* − '9 3 ∝ = '$%& /'Solvent 266

nabs:

Moles of gas absorbed

267

ninj:

Moles of gas injected

268

nSolvent: Moles of reactive component (amine) in the solvent

269

nEC:

Moles remaining in system overhead at equilibrium

270

∝:

CO2 loading in the solvent

271

Pinit:

Initial pressure (B: Batch Vessel; EC: Equilibrium Cell)

272

Pfinal: Final pressure (B: Batch Vessel; EC: Equilibrium Cell)

273

Vsys:

Volume (B: Batch Vessel; EC: Equilibrium Cell)

274

Vsol:

Volume occupied by the solvent (l: liquid; v: vapor)

(4)


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R:

Gas constant

276

T:

Temperature (B: Batch Vessel; EC: Equilibrium Cell)

277

Zinit:

Initial compressibility factor for PR-EOS (B: Batch Vessel; EC: Equilibrium Cell)

278

Zfinal: Final compressibility factor for PR-EOS (B: Batch Vessel; EC: Equilibrium Cell)

279

The Peng-Robinson equation of state (PR-EOS) was used to estimate the compressibility

280

coefficient to account for deviations from ideality. Under essentially all experimental conditions

281

relevant to post-combustion CO2 solvents, the compressibility coefficient ranges between 0.985

282

and 1.00, and is taken into account in the data analysis.

283

Results and Discussion

284

The most common reaction products formed from the reaction between CO2 and aqueous amine

285

solvents using weakly basic amines are amine carbamates and bicarbonates. In non-aqueous

286

solvent chemistry, carbonate ester products have been reported for strongly-basic compounds

287

such as amidines, guanidines, and phosphazine super-bases in protic solvents.26-31 We have used

288

nuclear magnetic resonance spectroscopy along with other indicators to investigate the product

289

of the reaction of Formulation 1 with CO2 via formation of amine carbamates. When fluorine

290

nuclei are present in alcoholic reactants, fluorine NMR is a convenient handle to distinguish

291

between the formations of carbamate or carbonate esters. The formation of a new product which

292

involves the alcohol will render new 19F resonances in the NMR spectrum. For Formulation 1,

293

19

294

1 depicts two possible reaction pathways along with the identification of the unique 19F signals

295

which appear in the 19F NMR spectrum. Figure 1 shows the changes observed in the 19F NMR

F NMR evidence indicates that the fluorinated alcohol is not involved in CO2 capture. Scheme


14

Page 15 of 31

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296

spectrum of the solvent before and after purging with CO2 (all fluorine spectra are proton

297

decoupled). If the fluorinated alcohol was involved in the capture, it is anticipated that five new

298

resonances should appear in the product corresponding to bonding of the conjugate base of the

299

alcohol to CO2. Instead, the four resonances for the chemically distinct fluorine nuclei of 2,2’,

300

3,3’, 4,4’, 5,5’-octafuoropentanol remain unchanged after exposure to CO2, and in fact only one

301

new resonance appears. The new resonance is the resonance for F1’ of the 2-

302

fluorophenethylamine carbamate, which upon complete conversion results in a spectrum with a

303

total of six unique

304

labeled F1, F1’, F2, F3, F4, and F5 shown in the lower left of Scheme 1 and labeled in Figure 1.

19

F resonances. The six unique resonances correspond to the fluorine nuclei

305 306

Scheme 1. Reaction pathways from 2-fluorophenethylamine and 2,2’, 3,3’, 4,4’, 5,5’-

307

octafuoropentanol starting materials. The reaction pathway on the left forms a carbamate salt and

308

has six unique resonances. The reaction pathway on the right forms a carbonate ester and has

309

five unique resonances


15


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Page 16 of 31

-

F1

+ NH3

O

F1

F2 F3 F4 F2 F5 F3 F H 4 F5

F3 F2

F4 F5

-116

-118

-120

-122

-124

-126

-128

-130

-132

HO

-134

F2 F3 F4 F2 F5 F3 F H 4 F5

-136

-138

+ NH3

-

O

+

F1

F3 F2

-140

NH O F1'

F1

F4

F1' F5

310

-116

-118

-120

-122

-124

-126

-128

-130

-132

-134

-136

-138

-140

311 312

Figure 1. Fluorine NMR spectra of 2-fluorophenethylamine and 2,2’, 3,3’, 4,4’, 5,5’-

313

octafuoropentanol before (top) and after (bottom) reaction with CO2 showing six unique

314

resonances in the product.

315 316

In the absence of any alcohol, 2-fluorophenethylamine reacts with CO2 to form a carbamate salt.

317

The reaction is shown in the supporting information, S1. The carbamate salt is a solid at room

318

temperature but is solubilized by many organic solvents. An NMR experiment was conducted in

319

deuterated chloroform solvent containing only 2-fluorophenethylamine. The 19F NMR spectrum

320

shows a single resonance for the F1 fluorine in 2-fluorophenethylamine. With formation of the

321

carbamate, a single new resonance appears (Supporting Information, Figure S2). The

322

corresponding 1H NMR spectra also indicate that the amine carbamate is formed. The 1H NMR


16

Page 17 of 31

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323

spectrum is shown in Error! Reference source not found.. Notably, the resonances for the

324

amide N-H and protonated –NH3+ become clearly observable upon reaction of the amine with

325

carbon dioxide (Supporting Information, Figure S3).

326

Examination of the 1H NMR spectrum of Formulation 1 (Figure 2) shows all of the anticipated

327

resonances for the two-component system prior to exposure to CO2, with overlap occurring

328

between H9’ N-H protons of the amine and H2’ methylene protons of the alcohol. Upon

329

exposure to CO2, new resonances are observed in the spectrum which correspond with formation

330

of 2-fluorophenethylamine carbamate. A resonance corresponding to the three H9’ protons of the

331

protonated nitrogen appear downfield just above the aromatic region and presumably this shift is

332

due to a hydrogen bonding effect from the protonated acidic alcohol. The H1 and H2 protons of

333

the fluorinated alcohol are unchanged and new methylene resonances H3 and H4 appear from the

334

phenethyl moeity of the carbamate. The signal for H8 appears between 4.5 and 5.0 PPM which is

335

characteristic of the amide N-H formed after reaction of 2-fluorophenethylamine with CO2 to

336

produce the carbamate. Monitoring the reaction by 13C NMR shows the emergence of new

337

resonances at 163 ppm, 40 ppm, and 30 ppm which are consistent with carbamate formation (see

338

supporting information, Figure S1).

339


17


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Page 18 of 31

340 341

Figure 2. Proton NMR spectra showing transformation of Formulation 1 into 2-

342

fluorophenethylamine carbamate.

343

The NMR experiments are conducted at a CO2 partial pressure of approximately 101 kPa. Below

344

this pressure the reaction products are anticipated to be the same as the species observed by

345

NMR, while at pressures above 101 kPa it may be possible to observe other species. The

346

presence of fluorine groups has been reported to improve CO2 solubility in physical solvents.32

347

Formation of carbonate esters could also occur at higher pressure. The partial pressure of carbon

348

dioxide in flue gas ranges from approximately 8 to 20 kPa.

349

The viscosity of the unreacted 2-FPEA/OFP system is 7.2 cP at an operating temperature of

350

40°C. Reacting 2-FPEA/OFP with CO2 leads to an increase in viscosity, presumably due to the

351

formation of a hydrogen bonding network. The viscosity of the CO2 saturated 2-FPEA/OFP at

352

40°C is 27.1 cP. Although there is an increase in the viscosity, it is relatively low compared to

353

other non-aqueous systems that commonly have viscosities > 100 cP after reaction with CO2.


18

Page 19 of 31

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354

The water solubility of Formulation 1 has been determined by Karl-Fischer titration. The lean

355

solvent, free of CO2, absorbs approximately three weight percent water at ambient temperature.

356

Above this threshold, water separates and the solvent becomes a two phase mixture. When the

357

solvent is fully saturated with CO2 at ambient temperature, it absorbs approximately nine weight

358

percent water before water separates as a second phase.

359 360

Vapor liquid equilibrium measurements of Formulation 1 were conducted from approximately

361

0.2 to 480 kPa to assess the overall CO2 working capacity. The results are shown in Figure 3

362

below. At 30°C, Formulation 1 absorbs CO2 at low partial pressures, measured down to

363

approximately 0.2 kPa where the loading is approximately 0.176 moles CO2/ mole amine. At

364

higher temperatures (80, 120 °C) Formulation 1 does not absorb a significant amount of CO2 at

365

a partial pressure of 0.2 kPa. At 120°C the solvent absorbs 0.137 moles CO2/ mole amine at 480

366

kPa.


19


1000

100

PCO2 (kPa)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 20 of 31

10 0

30 C 0 40 C 0 60 C 0 80 C 0 120 C

1

0.1 0.0

367

0.1

0.2

0.3

0.4

0.5

0.6

CO2 loading (molCO2/mol2FPEA)

368

Figure 3. Vapor liquid equilibrium isotherms for Formulation 1 at different temperatures.

369

The CO2 vapor liquid equilibrium measurements of Formulation 1 at 30°C shows a loading

370

higher than 0.5 moles CO2/ mole amine at pressures above 100 kPa. This can be explained by the

371

transition from chemical absorption to physical absorption. Based on the magnitude of Henry’s

372

constants for CO2 for the individual components at 30°C, it is within reason to postulate physical

373

dissolution of CO2 into the solvent as the cause of the CO2 capacity beyond the carbamate

374

theoretical maximum.33 Other amine-based systems which are primarily non-aqueous in

375

composition have also reported certain physical loading capacity.34

376

In aqueous amine systems such as a 30 wt% solution of monoethanolamine (MEA) in water, the

377

heat of absorption can be indicative of a change in reaction mechanism. With water as a reactant,


20

Page 21 of 31

378

the reaction pathway can include formation of bicarbonates and result in theoretical CO2

379

loadings as high as 1 mole CO2/ mole amine. A correlation between the magnitude of the heat of

380

absorption and the loading is observed which suggests that the carbamate product is formed at

381

loadings less than 0.5 mole CO2/ mole amine and generates a higher heat of absorption. Above

382

loadings of 0.5 mole CO2/ mole amine the carbamate is converted to the bicarbonate product

383

with less reaction heat. This correlation was observed for 30 wt% aqueous MEA and compared

384

to previous literature reports.5, 35 The measurements agreed well with previous reports and are

385

shown in Figure 4.

120

100

∆ Habs (kJ/molCO2)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


80

60

40

20 0.0

386

40oC (This work) o 80 C (This work) o 120 C (This work) o 40 C (Kim et al. 2014) o 80 C (Kim et al. 2014) o 120 C (Kim et al. 2014) 0.2

0.4

0.6

CO2 loading (molCO2/molMEA)

387

Figure 4. Heat of absorption measurement for 30 wt% aqueous MEA measured at 40, 80, and

388

120°C and compared to previously reported data.


21


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Page 22 of 31

389

The heat of absorption was measured for Formulation 1 and is shown in Figure 5 below. There

390

are several contrasts between the heat of absorption of Formulation 1 compared with 30 wt%

391

aqueous MEA. At 40°C (temperature anticipated for a post-combustion absorber application) the

392

reaction heat appears slightly higher in magnitude at loadings up to 0.25 mole CO2/ mole amine

393

(80.8-91.1 kJ/mole CO2) compared to that of MEA; however, above loadings of 0.25 mole CO2/

394

mole amine, the reaction heat of Formulation 1 decreases and has a value lower than that of

395

MEA. At a loading of 0.317 moles CO2/mole amine the heat was measured to be 70.6 kJ/mole

396

CO2 and decreased further to approximately 53.6 kJ/mole CO2 at 0.424 mole CO2/ mole amine.

397

Between 0.5-0.6 moles CO2/ mole amine there was a further decrease leading to values which

398

are similar to those observed for 30 wt% aqueous MEA at CO2 loadings higher than 0.6 mole

399

CO2/mole amine loading, i.e. beyond the carbamate regime.

400

Another contrast is that Formulation 1 shows lower heats of absorption at higher temperatures.

401

The authors speculate that this may be connected to changes in hydrogen bonding around the

402

carbamate at higher temperature or decrease in potential π-interactions of aromatic groups at

403

higher temperature. A change in the specific heat capacity of the solvent with loading must also

404

be considered as a factor in this observable. Testing of hypotheses to explain this aspect of

405

Formulation 1 and similar non-aqueous solvents is underway. This trend is opposite to that

406

observed for 30 wt% aqueous MEA, where the heat of absorption is measured to be higher as the

407

temperature increases. Thus, while the heat of absorption for 30 wt% aqueous MEA increases to

408

approximately 100 kJ/mol CO2 at 120°C, the heat of absorption for Formulation 1 at 120°C

409

decreases to approximately 50 kJ/mole CO2 and lower. At this temperature, the solvent loads a

410

relatively small amount of CO2 even at CO2 partial pressures approaching 500 kPa (Figure 3).


22

Page 23 of 31

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411

The heat of absorption is measured until a pressure of approximately 724 kPa is obtained in the

412

calorimeter.

413 414

Figure 5. Heat of absorption of Formulation 1 measured at 40, 80, 90, and 120°C.

415

The minimum thermal regeneration energy demand (i.e., reboiler heat duty) and optimal solvent

416

recirculation rate was estimated for Formulation 1 using three different modeling methods. A

417

validated “short-cut” modeling method developed and reported by Notz et al.36 was compared

418

with a Promax model for 30 wt% MEA and rate-based Aspen equilibrium Electrolyte-Non

419

Random Two Liquid Symmetric Reference (ENRTL-SR) model for 30 wt% MEA.


23


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Page 24 of 31

420

In the “short-cut” method, a simple equilibrium stage model based on a modified Kremser

421

equation is used to describe the absorber and regenerator. Equilibrium isotherms and estimated

422

caloric data including heat of vaporization, heat of absorption, and the heat capacity of the

423

solvent and water at absorber and regenerator temperatures are required as inputs to the method.

424

The reboiler duty was estimated by summing up the energy required for the sensible heat of the

425

solvent, the reaction heat of the solvent with CO2, the stripping heat of CO2 from the solvent, and

426

the reflux heat with a constant heat capacity of water averaged over the reflux and regenerator

427

outlet temperature.36 A Promax model was developed using the Electrolytic Extended Long

428

Range (ELR)- Soave-Redlich-Kwong (SRK) equation of state for representing equilibrium and

429

TSWEET Kinetics in the absorber and regenerator for describing reaction kinetics. This model

430

requires the user to provide column information so that the residence time in the column can be

431

calculated to fully model the reaction kinetics. The rate-based, Aspen equilibrium ENRTL-SR

432

model was developed in Aspen Plus V.8.6 for rate-based separations with a rigorous framework

433

to simulate the absorber and regenerator.37The thermodynamic properties were calculated with

434

ENRTL-SR to describe the liquid phase and Redlich-Kwong (RK) equation of state for vapor

435

phase. The model was validated for 30wt% MEA with open literature and in-house measured

436

data. The absorber model includes both equilibrium and kinetic rate-based controlled reactions,

437

whereas the stripper model comprises equilibrium rate-bases controlled reactions.

438

The minimum thermal regeneration energy estimates were compared to Formulation 1. Reboiler

439

heat duty estimates for Formulation 1 were performed using the “short-cut” method.

440

Formulation 1 can achieve a CO2 product pressure of 780 kPa at 120°C but was not considered

441

in short-cut method. Experimental VLE curves of Formulation 1 at absorber and regenerator


24

Page 25 of 31

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442

process conditions, heat of CO2 absorption, specific heat capacity, and vapor pressure of

443

Formulation 1 were used as inputs at the following process conditions:

444 445

•

Flue Gas Composition (mole %): N2: 66.90; O2: 2.35; H2O: 16.68; CO2: 13.26

446

•

Percent CO2 Captured: 90%

447

•

Temperature: Absorber: 40°C; Regenerator: 120°C

448

•

Crossover Exchanger Approach: 10°C

449

•

Pressure: Absorber: 101.3 kPa; Regenerator: 200 kPa

450


25


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Page 26 of 31

451 452

Figure 6. Predicted reboiler duty for NAS--CO2 capture system at different gas and liquid

453

flow rates

454

As shown in Figure 6, the three process models agree well for 30% MEA and the values agree

455

well with experimental reports. The two process models used for Formulation 1 do not agree as

456

closely, however both models show that Formulation 1 has the potential to reduce the thermal

457

regeneration energy requirement compared to conventional CO2 capture processes by 40%–50%.

458

Though the thermal regeneration energy savings is promising, fluorinated components are more

459

expensive than aqueous solvent components. For example, octafluoropentanol and 2-

460

fluorophenethylamine are three and thirty times more expensive than monoethanolamine,

461

respectively, based on current market offerings.

462

Conclusions


26

Page 27 of 31

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463

A non-aqueous solvent composed of 2-fluorophenethylamine and octafluoropentanol reacts with

464

CO2 to form an amine carbamate. Fluorine, proton, and carbon NMR show convincing evidence

465

for this reaction product. Measurements of the CO2 vapor-liquid equilibrium and heat of CO2

466

absorption show a potentially larger working capacity and lower heat of absorption at higher

467

temperatures as compared to MEA. Approximating process models indicate the potential to

468

lower reboiler heat duty in a commercial implementation.

469

Supporting Information(SI)

470

Supporting Information 1H NMR spectrum of Formulation 1 , 13C NMR spectrum of

471

Formulation 1, 19F NMR spectrum of 2-FPEA, reaction scheme of 2-FPEA carbamate

472

Acknowledgement

473

The authors would like to thank the United States Department of Energy - Office of Fossil

474

Energy (DE-FE0013865) and Research Triangle Institute for financial support for this work. The

475

authors would also like to acknowledge Ms. Kelly Amato at RTI for Karl-Fisher analysis, Dr.

476

Raghubir Gupta and Dr. Markus Lesemann for their respective executive roles in the research.

477 478

AUTHOR INFORMATION

479

Corresponding Author

480

*Marty Lail

481

Funding Sources

482

This work was funded by the United States Department of Energy, National Energy Technology

483

Lab, under award number DE-FE0013865.


27


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Page 28 of 31

484

ABBREVIATIONS

485

CCS, carbon capture and sequestration; MEA, monoethanolamine; AMP, 2-amino-2-methyl-

486

propanol; TEG, triethyleneglycol; NMR, nuclear magnetic resonance; RTIL, Room temperature

487

ionic liquids; IL, ionic liquid; 2-FPEA, 2-fluorophenethylamine; OFP, 2,2’, 3,3’, 4,4’, 5,5’-

488

octafuoropentanol; VLE, vapor-liquid equilibrium.

489

REFERENCES

490 491 492 493 494 495 496 497 498 499 500 501 502 503 504 505 506 507 508 509 510 511 512 513 514 515 516 517 518 519 520 521 522

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523 524 525 526 527 528 529 530 531 532 533 534 535 536 537 538 539 540 541 542 543 544 545 546 547 548 549 550 551 552 553 554 555 556 557 558 559 560 561 562 563 564 565 566 567


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30. Wang, C.; Mahurin, S. M.; Luo, H.; Baker, G. A.; Li, H.; Di, S., Reversible and robust CO2 capture by equimolar task-specific ionic liquid-superbase mixtures. Green Chem. 2010, 12, 870-874. 31. Luo, H.; Baker, G. A.; Lee, J. S.; Pagni, R. M.; Dai, S., Ultrastable Superbase-Derived Protic Ionic Liquids. J. Phys. Chem. B 2009, 113 (13), 4181-4183. 32. Gwinner, B.; Roizard, D.; Lapicque, F.; Favre, E., CO2 Capture in Flue Gas: Semiemperical Approach to Select a Potential Physical Solvent. Ind. Eng. Chem. Res. 2006, 45, 5044-5049. 33. Verseteeg, G. F.; Swaaij, W. P. M. v., Solubility and Diffusivity of Acid Gases (CO2, N2O) and Aqueous Alkanolamine Solutions. J. Chem. Eng. Data 1988, 33, 29-34. 34. Zong, L.; Chen, C.-C., Thermodynamic modeling of CO2 and H2S solubilities in aqueous DIPA solution, aqueous sulfalone-DIPA solution, and aqueous sulfalone-MDEA solution with electrolyte NRTL model. Fluid Phase Equilib. 2011, 306, 190-203. 35. Kim, I.; Svendsen, H. F., Heat of Absorption of Carbon Dioxide (CO2) in Monoethanolamine (MEA) and 2-(Aminoethyl)ethanolamine (AEEA) Solutions. Ind. Eng. Chem. Res. 2007, (46), 5803-5809. 36. Notz, R.; Tonnies, I.; Mangalapally, H. P.; Hoch, S.; Hasse, H., A short-cut method for assessing absorbents for post-combustion carbon dioxide capture. Int. J. Greenh. Gas Control 2011, 5 (3), 413-421. 37. Aspen Properties Reference Manual. Aspen Technology, Inc.: Burlington, MA, 2006.

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CO2 Capture Using Fluorinated Hydrophobic Solvents - Industrial

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