Dossier: Methodology for Process Development at IFP Energies nouvelles
Open Access
Issue
Oil Gas Sci. Technol. – Rev. IFP Energies nouvelles
Volume 71, Number 3, May–June 2016
Dossier: Methodology for Process Development at IFP Energies nouvelles
Article Number 40
Number of page(s) 21
DOI https://doi.org/10.2516/ogst/2015038
Published online 06 June 2016

© J. Magné-Drisch et al., published by IFP Energies nouvelles, 2016

Licence Creative Commons
This is an Open Access article distributed under the terms of the Creative Commons Attribution License (http://creativecommons.org/licenses/by/4.0), which permits unrestricted use, distribution, and reproduction in any medium, provided the original work is properly cited.

Notation

Ap: Particle surface area (m2)

ai: Thermodynamic parameter

bi: Thermodynamic parameter

T : Temperature (K)

b7: Water adsorption coefficient in reaction 7 (bar−1)

b8: Water adsorption coefficient in reaction 8 (bar−1)

: Gas concentration of compound i (mol/m3)

: Particle concentration of compound i (mol/m3)

: Gas axial dispersion coefficient (m2/s)

Deff, i: Effective diffusion coefficient (m2/s)

dp: Particle diameter (m)

Dm: Molecular diffusion coefficient (m2/s)

k7: Kinetic constant of reaction 7 (SI)

k8: Kinetic constant of reaction 8 (SI)

kgs: Gas-solid mass transfer coefficient (m/s)

Keq, i: i th thermodynamic constant

L: Particle characteristic length (m)

: CO2 partial pressure (bar)

PCOS: COS partial pressure (bar)

: CS2 partial pressure (bar)

: H2O partial pressure (bar)

: H2S partial pressure (bar)

Pt: Total pressure (bar)

Re: Reynolds number

r7: Reaction rate of reaction 7 (mol/s/kg cat)

r8: Reaction rate of reaction 8 (mol/s/kg cat)

rj: Reaction rate of reaction j (mol/s/kg cat)

R: Perfect gas constant (SI)

r: Radius coordinate (m)

Rp: Particle radius (m)

Sc: Schmidt number

Vp: Particle volume (m3)

Vsg: Superficial gas velocity (m/s)

z: Reactor elevation (m)

Greek symbols

εg: Gas holdup

εs: Solid holdup

εp: Particle porosity

δ: Film thickness (m)

μij: Stoichiometric coefficient

ρg: Gas density (kg/m3)

ρs: Solid density (kg/m3)

τ: Tortuosity

Introduction: the gas treatment chain and new challenges

Natural gases, when collected from the production wells, are commonly polluted with many contaminants. Among these contaminants are sulfur compounds and CO2. Throughout the oil and gas treatment chain, various steps intend to separate most of the undesired compounds from the profitable part of the natural gas. The desulfurisation of natural gas is usually performed in generic treatment processes and consists of removing H2S and/or CO2 to meet the export gas specifications.

Commercial gas specifications also focus on compounds other than H2S and CO2. New specifications have been imposed for many years to also remove most of other sulfur compounds from the natural gas. One of these, Carbonyl Sulfide (COS) is usually present in sour natural gases containing both H2S and CO2, in quantities which may reach 150 to 1 000 ppm vol., due to the exploitation of gas fields with an increase in acid gas compounds.

The market for natural gas is in constant increase, +3%/year for natural gas and +7%/year for Liquefied Natural Gas (LNG), and we can assume that there is a correlation between H2S content and COS content such that 70% of raw acid gas world reserves are affected by this problem.

In addition to the total sulfur specification on the treated gas, there are some specific constraints with COS:

  • after fractionation of the treated gas, 90% of residual COS will be recovered in the C3 fraction and 10% in the C2 fraction because of the respective boiling point (−50.2°C for COS and −44.5°C for propane) which could entail problems on downstream petrochemical units,

  • if the treated gas is aimed to be transformed into LNG, the liquefaction process does not admit COS.

For H2S and CO2 deacidification combined with COS constraint, different treatment processes can be envisaged. Hybrid or physical solvents present the advantage of achieving both H2S, CO2 and COS elimination in one unit. However, for a selective application, it is not possible to respect the required H2S/CO2 selectivity and hydrocarbon co-absorption leads to economically unattractive processes. In addition, some of these solvents are not stable at high temperatures. Regarding these disadvantages, chemical solvents are in most cases more attractive solutions.

Amine solvents are very often used for natural gas deacidification purposes as they can be adapted to various specifications and to a wide range of feed gas compositions.

When complete CO2 removal from the gas is required, for instance to achieve LNG plant feed specifications, formulated MDEA solvents are more frequently being used, replacing primary or secondary amines. MDEA, when it is used without activator, is a selective amine solvent. It targets H2S removal, but leaves some CO2 in the treated gas. This reduces the sweetening cost as some CO2 can be left in the treated gas, and increases H2S concentration in the produced acid gases to Sulfur Recovery Units (SRU). MDEA is also used for an additional acid gas enrichment step to prepare the feed load of SRU for an optimised design of the Claus unit.

When MDEA is used for a selective H2S removal from the gas, a reduced fraction of the CO2 is separated from the feed gas but an even lower fraction of the COS is removed. It is well known that MDEA has poor COS removal capabilities. Severe COS specification cannot be reached for gases with fairly high initial COS content even with higher solvent circulation rate and/or high number of mass transfer units in the contactor.

In other words, it is not possible to design a selective removal plant that removes COS and leaves amounts of CO2 in the treated gas, as a selective plant should aim to do. Therefore, the reduction of the sweetening cost by leaving some CO2 in the gas, so as to improve the quality of acid gas, is equally unachievable. The design will likely rely on the use of a primary or secondary amine or a formulated MDEA. If acid gas enrichment is needed to efficiently design the Claus unit, then a second acid gas enrichment unit using selective solvent is compulsory to improve the quality of acid gas and increase its H2S content.

The severe COS specifications in the export gas consequently affect the cost of gas treatment when this is carried out with amines solvents, especially when a selective H2S removal unit should help to minimise the number of units of the Sulfur Recovery Facilities or to reduce the size of the high pressure acid gas removal unit.

Regarding this process analysis and following a review of catalytic solutions, IFP Energies nouvelles (IFPEN) has developed the COSWEETTM process that combines a very deep COS removal through the complete hydrolysis of the COS in a catalytic reactor integrated within the selective removal of H2S by a MDEA amine unit. The COSWEETTM reactor can also be integrated with an Energized MDEA unit as it reduces the solvent flow rate to perform both acid gas removal and COS removal.

At the screening phase, IFPEN has selected a specific metal oxide based catalyst that is able to operate at moderate temperatures. Numerous data have been acquired in the laboratory on the thermodynamics and the kinetics of the hydrolysis reaction taking place on the catalyst. Based on these data, a specific prediction model (and an in-house simulator) has been developed to allow the precise and optimal design of the catalytic reactor. Details of these developments are presented in this paper.

The last part of the paper focuses on the interest of the COSWEETTM process when selective acid gas removal is required through a representative case study. The paper also discusses interest in the COSWEETTM process in terms of capital and operating expenditures (CAPEX and OPEX) compared to a formulated amine application.

1 Identification of the concept and process definition

1.1 Existing Sweetening Processes: Advantages and Disadvantages

The solvents based on alkanolamines are the most generally accepted and widely used of the many available solvents for removal of acid gas H2S and CO2 from natural gas streams. However, alkanolamine based processes do not provide technological solutions and attractive compromises between investment and operational costs when they target tight COS specifications from gas containing large amounts of COS [1]. When total CO2 removal is required, conventional amine processes can be adapted (using for instance primary and secondary amines or formulated MDEA), to remove COS down to very low levels in treated gas. This is generally achieved by increasing the solvent flow rate along with the absorber height.

Some processes have also secured the COS transformation into H2S and CO2 by hydrolysis in a hot zone of the amine absorber, as this is performed in a specific version of the HiLoadDEA process licensed by Prosernat [2].

Some competitors [3] propose the use of physical solvents in selective treating by using differences in equilibrium constants (i.e. solubility), the advantage being the good absorption of COS component but the two major drawbacks are:

  • hydrocarbon coabsorption which represents a loss of valuable hydrocarbons into the acid gas (lost sales of natural gas components (C1, C2) and lost sales of C3+, NGL, condensate;

  • required high partial pressure of acid gas.

Other competitors propose the use of hybrid solvents [4] to remove the COS and mercaptans and maintaining the selectivity. However, the application of hybrid solvents is adapted to acid gas enrichment units and these solvents have the same problem of hydrocarbon coabsorption as physical solvents.

MDEA is the attractive solution for maintaining the selectivity but for gases containing COS, the use of selective MDEA processes faces two difficulties. The reaction between MDEA and COS is low and MDEA is not prone for removing COS. Reaction mechanisms of COS with MDEA are the same as those of MDEA with CO2.

There is no direct reaction between MDEA and CO2. Absorption of CO2 is achieved through hydrolysis (Reaction 2 catalysed by MDEA. This reaction is very slow and this is the reason why MDEA solvent is selective:

Reaction 1

Reaction 2

As noted by Sharma [5], reactions between COS and MDEA are based on the same mechanism. COS undergoes hydrolysis to yield carbon dioxide and hydrogen sulfide [6] according to Reaction 3:

Reaction 3

Al-Ghawas et al. [7] proposed that tertiary amines act as base catalyst for the hydrolysis reaction of COS according to Reaction 4:

Reaction 4

Reaction 4 is even slower than the CO2 hydrolysis in Reaction 2.

The existing sweetening processes do not have the capability to remove all COS while maintaining high CO2/H2S selectivity. The presence of the COS in the sour gas is a major problem in the case of natural pipe gas specification. Most of the time, a total acid gas removal unit is designed to achieve COS elimination.

1.2 Evaluation of Different Schemes to Reach High COS Elimination

Based on its experience of catalyst development, IFPEN has investigated the possibility to remove the COS by hydrolysis (Reaction 5) through heterogeneous catalysis. Reaction 5 implies the removal of acid gases formed by the hydrolysis reaction:

Reaction 5

The philosophy of the development is to keep the high selective performance of MDEA solvent with a conventional acid gas (CO2 and H2S) absorption and achieve significant COS removal through a dedicated catalytic reactor. Both catalytic and process schemes were investigated to find the suitable configuration.

1.2.1 First Configuration: COS Hydrolysis Followed by Amine Treatment

The first configuration investigates the possibility to perform the hydrolysis upstream of the absorber as a pretreatment. Figure 1 represents the first process scheme in which feed gas is sent into a feed/effluent heat exchanger and enters to the reactor final a last heating step so as to reach the temperature required by the process. The COS free exhaust gas is sent to the absorber to remove acid gas.

thumbnail Figure 1.

Process scheme 1: COS hydrolysis followed by amine treatment.

The feed gas composition treated by the hydrolysis reactor contains the highest quantity of acid gas. A typical acid gas composition in natural gas (Tab. 1) was taken into account for the preliminary investigation of the process scheme.

Table 1.

Natural gas composition

The water content is 1 200 ppm vol. (saturation of the natural gas). Due to high acid gas content (COS and H2S), and as COS hydrolysis reaction is an exothermic equilibrium reaction, either low temperature and/or higher water content would be required to reach high COS conversion. This is however unacceptable as in these conditions operating temperatures required to reach high COS conversion would be below gas dew point. At process scale, this would lead to capillary condensation inside catalyst porosity. This is illustrated in Figure 2, which represents the COS conversion at thermodynamic equilibrium for different water contents. These calculations underline that is not possible to reach high COS conversion by adding a huge amount of water (10 vol.%).

thumbnail Figure 2.

Thermodynamic COS conversion versus temperature.

This explains the necessity to remove the CO2 and H2S content upstream the COS hydrolysis reactor due to the thermodynamic constraint.

1.2.2 Second Configuration: Amine Treatment Followed by COS Hydrolysis

The second configuration proposes to hydrolyse the COS downstream of the absorber contactor (Fig. 3). This configuration favours the thermodynamic equilibrium of the hydrolysis reaction because the absorber removes CO2 and H2S compounds by MDEA absorption.

thumbnail Figure 3.

Process scheme 2: amine treatment followed by COS hydrolysis.

The natural gas enters at the bottom of the absorber and flows upward through the column and is contacted at counter current with the lean solvent MDEA. At the top of the column the gas has reached CO2 and H2S specifications. The gas is heated to the operating temperature of the process (100-180°C) and is sent to the hydrolysis reactor. The exhaust gas contains H2S produced by the hydrolysis reaction which should be removed by an additional reactor step. The typical purification reactor to remove low H2S concentration (≈ 100 ppm vol.) from a gas is a zinc oxide fixed bed reactor. Reaction 6 takes place at higher temperature than the hydrolysis reaction and requires an additional heater:

Reaction 6

Additionally, a lead-lag configuration is mandatory because zinc oxide is non-regenerable.

Figure 4 represents the COS conversion at equilibrium as a function of the temperature with the acid gas composition at the inlet of the hydrolysis reactor. Contrary to the first configuration, there is no thermodynamic limitation to reaching high COS conversion.

thumbnail Figure 4.

Thermodynamic COS conversion versus temperature.

A preliminary design of the hydrolysis reactor and ZnO reactor was performed based on this configuration. A shortcut reactor model for the hydrolysis reactor was developed by IFPEN for the preliminary design based on a catalyst and a kinetic law from the literature [8]. The composition of the hydrolysis feed gas reactor is summarised in Table 2. The hydrolysis reactor design is presented in Table 3 and the ZnO reactor design is presented in Table 4.

Table 2.

Hydrolysis feed gas reactor

Table 3

Hydrolysis reactor design

Table 4.

ZnO reactor design

The shortcut model validates the feasibility of the hydrolysis reactor in term of performance and the design is acceptable.

The zinc oxide reactor was designed with another shortcut model. The results of the design show that the configuration is not suitable for the H2S removal. Indeed, the huge amount of solid is not economically acceptable and implies a considerable size of the reactor. In fact, the residual H2S content (several tens of ppm) to be treated downstream of the hydrolysis reactor is too important due to the huge natural gas flow rate. The CAPEX of hydrolysis section plus zinc oxide reactor and heat exchanger equipment could not be envisaged.

An alternative solution to remove H2S reformed by COS conversion should be envisaged. Other processes exist on the market but the associated costs are not acceptable. The suitable solution will be to re-introduce the exhaust gas of the hydrolysis reactor into the upper part of the amine absorber in order to remove the last trace of the H2S.

1.2.3 Third Configuration: Combined COS Hydrolysis and Amine Treatment

The third configuration (Fig. 5) proposes to sweeten the acid gas from the hydrolysis reactor inside the absorber. The feed natural gas enters at the bottom of the absorber column in order to remove the bulk CO2 and H2S acid gases. The objective is the same as configuration 2 (amine treatment followed by COS hydrolysis) which is to avoid thermodynamic limitation. In a second step, the gas is withdrawn from the absorber and sent to the catalytic reactor after two heating steps. At the outlet of the hydrolysis reactor, the gas no longer contains COS and is sent back to the upper part of the absorber. H2S formed by the COS conversion is sweetened by MDEA lean solvent in countercurrent.

thumbnail Figure 5.

Process scheme 3: Combined COS hydrolysis and amine treatment.

The advantage of this configuration is the use of the same absorber column to remove the acid gas to reach the natural gas specifications and to remove the last decades of H2S ppm vol. present at the outlet of the catalytic hydrolysis reactor. This configuration implies two dedicated sections of the absorber, due to the pressure drop of the heat exchanger and the hydrolysis reactor.

The preliminary design of this last process scheme, with a combined COS hydrolysis and amine treatment, presents the best performance and it is the preferable configuration.

1.3 Definition of the COSWEETTM Process

The COSWEETTM process (Fig. 6) is based on the preferred process scheme 3 (combined COS hydrolysis and amine treatment) and is presented in the final configuration.

thumbnail Figure 6.

COSWEETTM process scheme.

Most of the acid compounds from the sour gas are removed by the solvent in the lower section of the absorption column. The gas from this first bulk removing stage is heated to an optimised temperature before entering the catalytic reactor where the COS is hydrolysed into CO2 and H2S. The produced gas is introduced into the upper part of the absorption column to complete the H2S removal. Therefore the absorption column needs two solvent feeds. Because of the low acid gas content in the upper section, the solvent flow rate in this section can be strongly decreased, allowing a reduction in column diameter.

The COSWEETTM section is integrated to a very simple and conventional amine unit. The configuration can be adapted to any arrangement of Acid Gas Removal Unit (AGRU):

  • the solvent used in a COSWEETTM unit is the one produced by the regenerator of the AGRU unit, with no need for extra duty or modified configuration of the stripping section;

  • the absorber can be fitted with trays or with packing beds;

  • the hydrolysis of COS in a reactor destroys COS, and it almost completely reduces the COS content in the flash gas and in the acid gas.

Some special arrangements of the reactor section, especially some dedicated arrangements to knock out the liquids upstream of the reactor, mitigate the risk of any solvent carry-over which could drastically reduce the hydrolysis catalyst activity.

2 Process development and first techno-economic evaluation

2.1 Validation of the Concept: Catalyst Choice and First Techno-Economic Evaluation

The first techno-economic comparison (Tab. 5) is based on the COSWEETTM scheme versus a total removal absorption scheme by use of the solvent DEA (reference case).

Table 5.

First techno-economic evaluation

The COSWEETTM process allows to maintain 1.8% of CO2 in the treated gas by achieving a deep COS removal (1 ppm vol.). To reach 1 ppm vol. of COS in the treated gas with DEA process implies the total removal of CO2. The CAPEX reduction is significant even when considering the cost increase due to the additional reactor and heat exchanger. Indeed, the reduction of the size of the column and of the solvent flow rate balances the CAPEX increase of the COSWEETTM equipments.

2.2 Definition of the Process Phase Development

The preliminary study resulted in high interest in the COSWEETTM process for natural gas application to remove COS to meet very severe specifications (< 1 ppm vol.) allowing high H2S/CO2 selectivity to be kept. The process development approach could be illustrated by Figure 7.

thumbnail Figure 7.

Process phase development.

Referring to Figure 7, based on a commercial catalyst, a shortcut reactor model has been developed to identify the best process scheme to remove COS in parallel to the AGRU. The technical evaluation allowed the determination of the suitable scheme that could be economically investigated. If the economic aspect is acceptable, the scheme is protected with a registered patent and a more thorough study could start.

Catalyst laboratory tests should be performed to determine the activity of the catalyst in the process operating condition and a rigorous kinetic model should be developed. Based on the kinetic and process constraint a fixed bed reactor model was developed. Additional experimental lab tests were performed to validate the model at low and high conversion for several operating conditions.

Based on the preferred scheme a patent was registered to protect the COSWEETTM process [9] (Fig. 8).

thumbnail Figure 8.

COSWEETTM process scheme [9].

3 Experimental study

Following techno-economic evaluation, an experimental study has been carried out in order to obtain data on catalyst activity toward the COS hydrolysis reaction. Objectives were primarily to validate the feasibility of COS conversion reaction under the industrial COSWEETTM process conditions. Experiments were also performed to obtain information on reaction kinetics to determine the critical parameters that may affect COS conversion and to get data to be used to develop a kinetic modeling tool. From these experiments, the dependence of reaction kinetics toward a set of parameters that have to be considered for the COSWEETTM process design has been underlined.

3.1 The Pilot Plant

A schematic representation of the experimental set-up used for kinetics measurements is reported in Figure 9.

thumbnail Figure 9.

Experimental equipment for kinetic measurements.

This equipment can be divided into three sections, as represented in Figure 9:

  1. a feed preparation zone where the different gas are mixed to build the feed gas. COS, CO2, He and H2S are supplied from gas tanks with specific gas compositions provided by Air Liquide. An helium and water mixture is prepared using a water saturator set-up where a helium flow bubbles in water heated at a controlled temperature. This controlled water saturated helium flow is mixed to the mixture of dry COS-H2S-CO2-He gas to reach the desired water content;

  2. a reaction zone, which basically consists of a cylindrical fixed bed reactor filled with the COS hydrolysis catalyst and heated at the desired temperature;

  3. an analytic set-up to analyse and quantify the reactions products through mass spectrometry. Initial COS gas contents and COS gas contents downstream of the hydrolysis reaction zone are measured to determine COS conversion rates as a function of the operating conditions.

3.2 Experimental Results

COS hydrolysis is thermodynamically very favoured at low temperature as shown in Figure 10. Preliminary long residence time tests have been first carried out to measure the occurrence of any inhibition effect or lack of catalyst activity under the industrial COSWEETTM process conditions. COS conversion rates have been measured for different temperatures from 100 to 200°C. These experiments show that the selected catalyst exhibited highly satisfactory activity toward the operating conditions imposed as it allowed nearly complete COS conversion close to thermodynamic equilibrium to be reached.

thumbnail Figure 10.

Free enthalpy and equilibrium constant of the COS hydrolysis reaction as a function of temperature (thermodynamic data obtained from HSC Chemistry v6.1 [10]).

Lower residence time tests have been performed to evaluate the kinetics of COS conversion as a function of experimental parameters: temperature, GHSV (Gas Hourly Space Velocity) or residence time, pressures H2O, COS, H2S, CO2. The operating condition ranges explored are summarised in Table 6. These data were used to produce a design model of the COS hydrolysis reactor with the selected catalyst.

Table 6.

Experimental range of acid gas partial pressures, temperature and space velocity. GHSV stands for Gas Hourly Space Velocity. NTP for Normal Temperature and Pressure

It was shown that the moderate presence of H2S and CO2 had no significant impact on the kinetics of the COS hydrolysis reaction in the operating range tested. At higher conversions, the presence of H2S and CO2 can decrease the conversion performances due to thermodynamic equilibrium.

The addition of water has a moderate impact on the conversion of COS, as water is already in large excess compared with the expected COS levels. It was verified that for industrial residence time ranges, the hydrolysis reaction of COS was not inhibited, and that the catalyst activity was not affected by the presence of water.

Experiments have also been performed on crushed and uncrushed catalyst particles to take into account potential macroporous diffusion limitations inside catalyst particles. Indeed, significant transport limitations have been highlighted from these experiments, due to the high reaction rate. In agreement with experimental observations, these transport limitations are also demonstrated by the modeling performed for both lab-scale experiments and industrial case study (Sect. 4).

Thiele modulus calculations have been performed in order to have an estimation of the potential limitations encountered, which are function of the relative importance of reaction kinetic towards mass transport [11]. Thiele modulus ϕ is given by the following relation:

with k the rate constant as calculated from the modeling (Sect. 4), [COS] the COS concentration, n the order of reaction (= 1 in the modeling), D the COS effective molecular diffusion coefficient (calculated from the Füller correlation, Eq. 11), L c the particle diameter.

Thiele modulus calculations for some operating conditions for lab scale experiments as a function of particle size (crushed versus uncrushed particles), and for a typical industrial case are given in Table 7. This shows in every case the occurrence of mass transport limitations inside catalyst porosity, due to high catalyst activity (high kinetic rate). The extent of mass transport limitations are more pronounced with increasing particle size and/or pressure (the latter affecting molecular diffusivity). As a consequence, diffusion limitations are more pronounced at an industrial scale.

Table 7.

Estimation of transport limitations from Thiele modulus calculations for some experimental conditions lab scale versus industrial scale

Collected kinetics data as a function of reactor size, gas residence time, temperature and reactant partial pressures were then used to develop a reactor modeling tool taking into account all of the relevant operating parameters that may affect reaction kinetics and process performances.

4 COSWEETTM reactor model and simulator

An industrial reactor model is developed for scale-up purposes. This model is based on a kinetic model validated with experiments obtained in a lab scale fixed bed reactor. First of all, the lab scale reactor is described taking into account all the limitations (external mass transfer and intra particle diffusion) in order to catch the so-called intrinsic kinetic parameters for COS and CS2 elimination reactions. Then, this kinetic model is implemented in a complete industrial reactor model taking into account potential limitations. Then, it is used for scale-up and optimisation purposes.

4.1 Industrial Reactor Configuration

The industrial reactor is a two-phase (gas-solid) fixed bed system. It works under isothermal conditions due to the low amount of impurities to be removed. The pressure drop along the reactor should be below 1 bar (industrial specification). Different catalyst shapes can be used (spheres or cylinders) with an average size ranging from 1 to 3 mm. Typical operating conditions are in the range of 30-100 bar for the pressure, and 100-200°C for the temperature.

4.2 Gas Composition and Impurities to Be Removed

This process is used for natural gas treatment. A typical gas composition is given in Table 8.

Table 8.

Typical composition of natural gas

COS and CS2 are the main impurities to be removed. The final concentration of both species should be lower than 1 ppm mol.

4.3 Reaction Scheme, Thermodynamics and Kinetics

As mentioned previously, COS and CS2 can react with water in order to form H2S, COS and CO2 compounds according to the following reactions:

Reaction 7

Reaction 8

Both reactions are reversible. Thermodynamic equilibrium constants depend on the temperature as shown in Equation (1). Table 9 gives the corresponding thermodynamic parameters for the equilibrium constants:(1)

Table 9.

Equilibrium constants

For both reactions, an Eley-Rideal mechanism was used in order to take into account the inhibition effect of water. Kinetic rate expressions are given in Equation (2):(2)

Kinetic parameters for reaction 7 were estimated from lab scale experiments (Tab. 10). For reaction 8, kinetic parameters are taken from the literature [12].

Table 10.

Kinetic parameters

4.4 Reactor Modeling

Material balances are written for each compound at different scales: in the gas flow, in the external mass transfer film around the catalyst pellet, and inside the catalyst porous network. For the gas flow, a dispersed-plug flow model was used to take into account a potential back-mixing effect. Equation (3) gives the corresponding gas material balance:(3)

Gas axial dispersion coefficient was estimated using the Gunn correlation [13] (Eq. 4):

with

and (4)

The material balance in the external film is given by Equation (5):(5)

with (film thickness)

and (characteristic length of the catalyst)

The mass transfer coefficient is given by the Yoshida correlation [14] which takes into account the gas flow pattern around the particle (Eq. 6):(6)

Gas material balances should respect the equation of state . Then, after summing all the gas equations and introducing the equation of state, we obtain the corresponding equation for the gas velocity:(7)

Particle material balances (Eq. 8 and 9) are written for two particle shapes (cylinders and spheres):(8) (9)

with εp = particle porosity.

The effective diffusion coefficient is a function of the molecular diffusion coefficient of each compound, the porosity and the tortuosity of the catalyst (as shown in Eq. 10):(10)

Molecular diffusion coefficients are estimated using the Füller correlation [15] (Eq. 11):(11)

The effects of pressure and temperature were taken into account for the molecular diffusion coefficient.

Reactor pressure drop is calculated using the Ergun equation or the Handley equation according to the gas flow regime (Eq. 12):

Ergun:

Handley et al.:(12)

4.5 Numerical Aspects

A spatial discretisation of the partial derivative equations was done using an upwind finite differences scheme for the convection terms and a centered finite differences scheme for the diffusion or dispersion terms. For time integration, the Lsode solver was used based on a predictor-corrector algorithm [16]. An excel interface coupled with Matlab (for 3D visualisation) was developed.

4.6 COSWEETTM Simulator

Figure 11 shows the model interface. For one simulation, the reactor geometry, operating conditions, the catalyst geometry and the gas composition can be changed.

thumbnail Figure 11.

Model interface.

Figure 12 shows the COS intra-particle concentration profile at different reactor elevations.

thumbnail Figure 12.

Particle COS concentration profiles at different axial positions in the reactor.

As discussed in Section 3, transport limitations are indeed highlighted by the modeling performed for both lab-scale experiments and industrial case study. Significant transport limitations are observed at the bed inlet due to the high reaction rate. At the reactor outlet, there are less limitations due to lower COS gas content.

In both industrial and lab-scale, heat transfers have been neglected (isothermal conditions), as COS concentrations in every cases remain very low (<< 1%v).

4.7 Model Validation

The model was validated with experiments carried out in a lab scale fixed bed reactor for non-crushed and crushed catalysts. Figure 13 shows a comparison between experimental and modeling data for COS conversion. A good agreement was obtained.

thumbnail Figure 13.

Parity plot for COS conversion with crushed and non-crushed catalyst.

5 Techno-economic evaluation

5.1 Case Study 1: Deep COS Removal with Bulk CO2 Removal

The following example presents a CO2 bulk removal application. The solvent used for acid gas removal is a formulated MDEA which is effective in absorbing H2S and CO2 and some of the COS. The COSWEETTM process allows the removal of COS down to very low levels. The treated gas has to meet typical pipeline gas specifications: less than 4 ppm vol. H2S and less than 2.0 vol.% CO2. The COS specification is less than 0.5 ppm vol. The feed gas composition and conditions are given in Table 11.

Table 11.

Feed gas description

The process scheme is as described in Figure 14. The absorber operates at relatively high temperature (85-90°C) because there is constant and intensive reaction of CO2 with amine occurring on every stage of the column. The need for heating the extracted gas prior to entering the hydrolysis reactor is minimised, and no gas/gas exchanger is necessary, reducing the overall pressure drop.

thumbnail Figure 14.

Process scheme.

If a traditional formulated MDEA process can easily reach the CO2 and H2S specifications, it does not allow complete removal of COS, even if the initial content is very low. Destruction of the COS by hydrolysis in a COSWEETTM catalytic section allows slipping of the required amount of CO2 with the treated gas, substantially reducing the size and the energy consumption of the amine section that would be traditionally designed with formulated MDEA. Table 12 presents the main sizing data of the COSWEETTM process which guarantees the three required specifications to be met.

Table 12.

Main sizing data - COSWEETTM

The extra heat duty needed for the removal of COS is less than 6% of the amine solution regenerator reboiling duty. The pressure drop of the COSWEETTM hydrolysis section is less than 1 bar.

5.2 Case Study 2: Deep COS Removal with Selective Sweetening

The main purpose of the COSWEETTM process is the removal of COS with the use of selective gas sweetening schemes using amine solvent that can produce a rich acid gas to the SRU, achieving at the same time a severe COS specification on the sweetened gas. The purpose of this case study is to explore the benefits of selective sweetening schemes, in addition to the removal of COS in high pressure AGRU.

The raw gas is first sent to the Acid Gas Removal Unit (amine based AGRU) and the separated acid gases are sent to a claus SRU followed by a selective amine Tail Gas Treatment Unit (TGTU). The overall sulfur recovery (in term of H2S conversion) must be higher than 99.9%. The case study, based on different AGRU schemes that remove COS (COSWEETTM or formulated MDEA) focuses on the impact of the different AGRU designs on the design of the sulfur recovery chain.

The design bases are presented hereafter. The feed gas composition on a wet basis along with conditions is given in Table 13.

Table 13.

Feed gas composition and conditions

The treated gas has to meet typical pipeline gas specifications for H2S and CO2: less than 4 ppm vol. H2S and less than 2.0 vol.% CO2. The COS specification is less than 1 ppm vol. The incinerator effluent specifications impose an effective destruction of all hydrocarbons, a reduced amount of sulfur compounds, and low sulfur emissions. Treated gas and incinerator effluent specifications are summarised in Table 14.

Table 14.

Treated gas and incinerator effluent specifications

Two treatment options are considered:

  • in base treatment case (Fig. 15), sweetening of the gas is achieved with formulated MDEA. This unit is designed to achieve the COS specification: the solvent flow rate and number of stages in the absorber are increased to enhance COS removal. This results in an almost complete co-absorption of CO2 from the gas. As a consequence the H2S content of the acid gas feeding the Claus unit is low, and the required overall sulfur recovery cannot be achieved. An Acid Gas Enrichment Unit (AGEU), using selective MDEA (MDEAmax process), is installed prior to the sulfur recovery chain to produce an acid gas of adequate quality to achieve the targeted sulfur recovery efficiency;

    thumbnail Figure 15.

    Base treatment case: formulated MDEA and AGEU.

  • in the second treatment option (Fig. 16), sweetening of the gas is done with a COSWEETTM process combined with the use of MDEA as amine solvent. This process achieves the severe COS specification and a sufficient CO2 slippage with the treated gas to eliminate the need for an AGEU prior to sulfur recovery.

    thumbnail Figure 16.

    Selective AGRU + COSWEETTM.

Units design and performances are presented in Tables 14 to 17. Significant sizing data and performances of the AGRU for the 2 options are summarised in Table 15.

Table 15.

AGRU - sizing data and performances

Main performance data of the sulfur recovery facilities are presented in Table 16.

Table 16.

Sulfur recovery - sizing data and performances

The capital costs of the treatment units for the 2 options are given in Table 17.

Table 17.

CAPEX comparison

Table 18 indicates the yearly cost of steam and power, in the same arbitrary unit as used in Table 17.

Table 18.

OPEX comparison

These results demonstrate that COSWEETTM offers an attractive solution for deep COS removal when selective H2S removal would be desired. It reduces the necessary amine solution circulation rate and simplifies the sulfur recovery facilities. CAPEX and OPEX savings of the COSWEETTM/selective amine integrated scheme compared to the usual formulated MDEA scheme are quite substantial (more than 7% for the CAPEX and more than 65% for the OPEX). This is due to the high efficiency of the hydrolysis catalyst at low temperature and the fact that while amine circulation rate in the AGRU is maintained at a low level, the H2S selectivity is increased without the need of an AGEU. This low temperature catalyst allows a good heat integration of the COS removal section minimising energy consumption.

Conclusion

The COSWEETTM process development is an example of the result of teamwork that benefits from the skill set of IFPEN, combined with an industrial approach integrating industrial constraints.

After a complete analysis of the gas treatment market, the gas treatment chain, existing processes and new challenges regarding raw acid gas compositions – with increasing amounts of acid gas contents like H2S and CO2 but also COS which contribute to the total sulfided compounds – an important challenge has been identified: what solution can be developed for COS removal while keeping a certain amount of CO2 to meet gas pipe specifications?

A complete internal review of existing technical solutions for removal of COS, combined with other technical and industrial constraints such as hydrocarbon solubility, has been performed in order to select the catalytic solution to achieve deep COS specification. The best catalyst has been selected regarding its performance at low temperatures in order to optimise the integration in an industrial gas treatment chain.

Then different process schemes have been identified and short-cut models have been developed in order to check the technical capability of these different process schemes, taking into account gas compositions and thermodynamic constraints, but also the economic evaluation of each scheme, in order to check the economic viability of each solution. Based on this study, the COSWEETTM scheme has been identified, a reactor model has been developed integrating thermodynamic and kinetic models, and kinetic measurements have been performed in order to collect all required data to estimate the model parameters.

COSWEETTM is a process developed for the treatment of natural gases containing COS. It is based on a combination of deacidification with any alkanolamine solution and of COS hydrolysis on a metal oxide based catalyst.

A characteristic of the COSWEETTM process is that it achieves almost complete COS hydrolysis at moderate temperature, making the process more attractive due to substantial heat transfer savings from lower heating of the gas to hydrolysis temperature and reduced cooling of the hydrolysed gas back to amine absorber temperature.

The COSWEETTM process can be combined with any amine, such as MDEA or formulated MDEA. This allows removal of COS down to very low specifications while maintaining some CO2 in the treated gas. When sulfur recovery is associated with the gas sweetening, using COSWEETTM to convert COS allows maintenance of a high H2S/CO2 ratio in the acid gas, which is required to optimise the gas treatment chain, including SRU. COSWEETTM removes COS in the gas phase, meeting strict COS specification without an increase in amine solvent flow rate and reboiler duty. Evaluations have been performed on CAPEX and OPEX and they show that COSWEETTM allows a reduction in the investment cost as well as the operating cost when compared to a formulated MDEA unit sized to reach the same COS specification.

When selectivity is required on a gas treatment, COSWEETTM provides an attractive and efficient solution as it improves the design of selective amine plant, while still ensuring the quality of the acid gas. Significant savings can be expected for the high pressure AGRU and also at the SRU since the SRU feed has an increased H2S content and low levels of hydrocarbons contaminants, without needing a dedicated acid gas enrichment unit.

References

  • Kohl A.L., Nielsen R.B. (1997) Gas Purification, 5th Edition, Houston, Gulf Publishing Company.
  • Cadours R., Magné-Drisch J., Normand L., Roquet D., Perdu G. (2006) COS removal from natural gases by absorption in alkanolamine solutions, 85th Annual GPA Convention, Grapevine, Texas, Marsh 5-8.
  • Korens N., Simbeck D.R., Wilhelm D.J. (2002) Process Screening Analysis of Alternative Gas Treating and Sulfur Removal for Gasification, Prepared for U.S. Department of Energy by SFA Pacific, Inc., Revised Final Report.
  • Torabi Angaji M., Ghanbarabadi H., Karimi Zad Gohari F. (2013) Optimizations of Sulfolane concentration in propose Sulfinol-M solvent instead of MDEA solvent in the refineries of Sarakhs, Journal of Natural Gas Science and Engineering 15, 22–26. [CrossRef]
  • Sharma M.M. (1965) Kinetics of reactions of carbonyl sulphide and carbon dioxide with amines and catalysis by Brönsted bases of the hydrolysis of COS, Transactions of the Faraday Society 61, 681–688. [CrossRef]
  • Svoronos P.D.N., Bruno T.J. (2002) Carbonyl sulfide: A Review of its chemistry and properties, Industrial & Engineering Chemistry Research 41, 5321–5336. [CrossRef]
  • Al-Ghawas H.A., Ruiz-Ibanez G., Sandall O.C. (1989) Absorption of carbonyl sulfide in aqueous methyldiethanolamine, Chemical Engineering Science 44, 3 631–639. [CrossRef]
  • Tong S., Dalla Lana I.G., Chuang K.T. (1993) Kinetic modeling of the hydrolysis of carbonyl sulfide catalyzed by either titania or alumina, The Canadian Journal of Chemical Engineering 71, 392–400. [CrossRef]
  • Gazarian J. (2009) Gas deacidizing method using an absorbent solution with COS removal through hydrolysis, Patent application EP2480315.
  • HSC Chemistry, version 6.1 (2007) Outotec Research.
  • Thiele E.W. (1939) Relation between catalytic activity and size of particle, Industrial & Engineering Chemistry 31, 7 916–920. [CrossRef]
  • Tong S., Dalla Lana I.G., Chuang K.T. (1995) Kinetic modeling of the hydrolysis of carbon disulfide catalyzed by either titania or alumina, The Canadian Journal of Chemical Engineering 73, 2 220–227. [CrossRef]
  • Gunn D.J. (1987) Axial and radial dispersion in fixed beds, Chemical Engineering Science 42, 2 363–373. [CrossRef]
  • Yoshida F., Koyanagi T. (1962) Mass transfer and effective interfacial areas in packed columns, AIChE Journal 8, 3 309–316. [CrossRef]
  • Fuller E.N., Giddings J.C. (1965) A comparison of methods for predicting gaseous diffusion coefficients, J. Chromatogr. Sci. 3, 7 222–227. [CrossRef]
  • Radhakrishnan K., Hindmarsh A.C. (1993) Description and use of LSODE - the Livermore Solver for Ordinary Differential Equations, NASA Reference Publication 1327, Lawrence Livermore National Laboratory Report UCRL-D-113855.

Cite this article as: J. Magné-Drisch, J. Gazarian, S. Gonnard, J.-M. Schweitzer, D. Chiche, G. Laborie and G. Perdu (2016). COSWEETTM: A New Process to Reach Very High COS Specification on Natural Gas Treatment Combined with Selective H2S Removal, Oil Gas Sci. Technol 71, 40.

All Tables

Table 1.

Natural gas composition

Table 2.

Hydrolysis feed gas reactor

Table 3

Hydrolysis reactor design

Table 4.

ZnO reactor design

Table 5.

First techno-economic evaluation

Table 6.

Experimental range of acid gas partial pressures, temperature and space velocity. GHSV stands for Gas Hourly Space Velocity. NTP for Normal Temperature and Pressure

Table 7.

Estimation of transport limitations from Thiele modulus calculations for some experimental conditions lab scale versus industrial scale

Table 8.

Typical composition of natural gas

Table 9.

Equilibrium constants

Table 10.

Kinetic parameters

Table 11.

Feed gas description

Table 12.

Main sizing data - COSWEETTM

Table 13.

Feed gas composition and conditions

Table 14.

Treated gas and incinerator effluent specifications

Table 15.

AGRU - sizing data and performances

Table 16.

Sulfur recovery - sizing data and performances

Table 17.

CAPEX comparison

Table 18.

OPEX comparison

All Figures

thumbnail Figure 1.

Process scheme 1: COS hydrolysis followed by amine treatment.

In the text
thumbnail Figure 2.

Thermodynamic COS conversion versus temperature.

In the text
thumbnail Figure 3.

Process scheme 2: amine treatment followed by COS hydrolysis.

In the text
thumbnail Figure 4.

Thermodynamic COS conversion versus temperature.

In the text
thumbnail Figure 5.

Process scheme 3: Combined COS hydrolysis and amine treatment.

In the text
thumbnail Figure 6.

COSWEETTM process scheme.

In the text
thumbnail Figure 7.

Process phase development.

In the text
thumbnail Figure 8.

COSWEETTM process scheme [9].

In the text
thumbnail Figure 9.

Experimental equipment for kinetic measurements.

In the text
thumbnail Figure 10.

Free enthalpy and equilibrium constant of the COS hydrolysis reaction as a function of temperature (thermodynamic data obtained from HSC Chemistry v6.1 [10]).

In the text
thumbnail Figure 11.

Model interface.

In the text
thumbnail Figure 12.

Particle COS concentration profiles at different axial positions in the reactor.

In the text
thumbnail Figure 13.

Parity plot for COS conversion with crushed and non-crushed catalyst.

In the text
thumbnail Figure 14.

Process scheme.

In the text
thumbnail Figure 15.

Base treatment case: formulated MDEA and AGEU.

In the text
thumbnail Figure 16.

Selective AGRU + COSWEETTM.

In the text

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