Development and design of the in-situ regeneration section of Vitrisol®, a novel, highly
selective desulphurization process
Wermink, Wouter Nicolaas
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Publication date: 2019
Link to publication in University of Groningen/UMCG research database
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Wermink, W. N. (2019). Development and design of the in-situ regeneration section of Vitrisol®, a novel, highly selective desulphurization process. Rijksuniversiteit Groningen.
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121
Chapter 7: Vitrisol® a 100% selective process for H
2S removal in the presence
of CO
2Reproduced with permission from J. Nat. Gas Eng. 2017, 2, 1, 50-83. Copyright 2017 Scrivener Publishing.
Abstract
Procede developed a novel desulphurization process that is 100% selective for H2S removal from industrial gases in the presence of CO2, called Vitrisol®. The Vitrisol® process can be described with the following overall reaction equation:
( )
The performance of Vitrisol® is demonstrated for two typical applications in shale gas production by comparing them to a standard amine treating process. The remainder toxic, acid gas produced by the latter technology is compressed for acid gas injection.
From the results it can be concluded that significant reductions in OPEX can be achieved by using Vitrisol® as depicted in the energy consumptions of the overall process. Contrary to the amine process, Vitrisol® does not require additional treatment of the off-gas stream as the H2S is directly converted to crystalline sulphur and CO2 that can be emitted to the environment.
This study illustrates clearly that it is advantageous to first remove H2S from a gas stream containing both H2S and CO2 prior to CO2 removal to reduce OPEX.
7.1 Introduction
Hydrogen sulphide (H2S) is a highly toxic and corrosive gas. Removal of H2S from acidic gas streams, such as natural gas, industrial gas or biogas, is important for safety, health, environmental and economic reasons. Several regenerative and non-regenerative H2S removal processes are readily available, which are economically viable only for specific gas compositions and gas flow rates. Apart from non-regenerative H2S removal by the use of e.g. adsorbents, all the regenerative aqueous liquid redox desulphurization processes (e.g. THIOPAQ, LO-CAT, SulFerox) capture CO2 to varying extents besides H2S.
The conventional method of removing H2S from natural gas is using an amine process. Subsequently, the H2S in the stripper gas is converted to elemental sulphur by a consecutive Claus process. For natural gas fields, usually containing more CO2 than H2S, this will result in an inlet acid gas stream for the Claus process that is low in H2S and high in CO2 content. The inlet gas stream should contain at least 20 mol.% of H2S to be able to produce a stable flame in a Claus furnace. Modification of the Claus process is needed between 20 mol.% and 50 mol.% H2S in the inlet acid gas stream. Above 50 mol.% H2S content no modification of the Claus process is required.1,2 Moreover, owing to the coabsorption of CO2 the regeneration costs of the amine process are substantially increased.
The novel Vitrisol® desulphurization process3 is based on the removal of H2S by precipitation with copper sulphate (CuSO4) in an aqueous, acidic solution. Copper sulphide (CuS) and sulphuric acid are formed in the gas treating process:4,5
( ) ( ) (7.1)
The Vitrisol® process is able to remove H2S from acidic gas streams without the coabsorption of CO2.4,6 Because the precipitation reaction occurs rapidly, the removal of H2S is limited by mass transfer in the gas phase. The Vitrisol® process is able to remove in one stage more than 99.9+% of
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the H2S present in the gas phase and has the typical characteristics of very selective H2S scavengers. For the absorption process required for the removal of H2S, there is no real process conditional constraint and the operating pressures and temperatures can vary at least between 0.1-10 MPa and 283-363 K, respectively.
A Vitrisol® pilot absorber was built to remove H2S from biogas, obtain representative samples of CuS and to verify design rules. Operational boundary conditions were determined with respect to continuous operation in the absorber and batch-wise operation of the absorption liquid.
The current status of the Vitrisol® process is scavenger-like application. Cu2+, the active compound in the absorption liquid, becomes depleted during H2S removal. It must be noted, however, that nowadays copper is an expensive commodity; therefore increasing amounts of H2S lead to increasing operational costs. To reduce the operational costs for large amounts of H2S and/or large scale applications, a regeneration step was developed to replenish Cu2+. The regeneration step is based on an operation encountered in copper ore processing, i.e. the dissolution of CuS with ferric sulphate (Fe2(SO4)3).7,8 Copper sulphate, elemental sulphur (S°) and ferrous sulphate (FeSO4) are produced in this process:
( ) ( ) ( ) (7.2)
Ferrous sulphate can be re-oxidized to ferric sulphate with O2 according to:
( ) (7.3)
Resulting in the overall net reaction for the removal of H2S:
( ) (7.4)
For the development of the regeneration process, the reaction behaviour of the parallel reactions occurring during the dissolution of CuS, i.e. Reactions 7.2 and 7.3, respectively, were investigated.
Wermink and Versteeg6,9 studied the oxidation of ferrous ions in acidic sulphate solutions (Reaction 7.3), and proposed kinetic equations derived by using both the data obtained for the initial reaction rates and the experimentally determined Fe2+ concentration profiles, respectively.
Furthermore, Wermink and Versteeg10 investigated the behaviour of the oxidation of ferrous ions in acidic sulphate solution, in the presence of Cu2+. It was concluded that Cu2+ enhanced the oxidation rate of Fe2+, however, some of the experiments were affected by the rate of mass transfer of oxygen. Besides Fe2+ oxidation, Wermink and Versteeg6,11 studied the dissolution reaction of CuS with Fe3+ (Reaction 7.2). Representative samples of CuS, obtained from Vitrisol® pilot absorber operations,12 were used in the study. It was concluded that an increase in temperature increased the rate of dissolution. Full conversion of CuS could be obtained, independent of temperature.
From the above mentioned investigations it was concluded that relatively mild conditions are required for the regeneration process, i.e. temperatures below 373 K and at pressures ranging from atmospheric to 1 MPa. No need of vast amounts of energy are required for the regeneration of the solvent, therefore this process has an extremely low energy footprint. Furthermore, the Vitrisol® liquid can be completely regenerated resulting in a solvent with fully restored activity and crystalline sulphur. The conditions required in the regeneration process are case dependent, e.g. on the amount of H2S to be removed.
To demonstrate the applicability of the Vitrisol® technology, two (conceptual) process designs of the Vitrisol® process have been compared to standard amine treating processes of cases previously
123 published by Weiland and Hatcher.13 Acid gas injection was chosen as downstream processing of the toxic, acid strip gas that is obtained as outlet stream from the amine treating process.
7.2 Case definition
The cases used to evaluate the process designs of the Vitrisol® process and standard amine treating processes are examples of shale gas from British Columbia and an example of one of the gas plants built to process gas from fields in the Barnett shale, as previously published by Weiland and Hatcher13 (see Table 7.1). In the simulations of the present study, the gases were considered to be saturated with water.
Table 7.1: Case specifications.
Gas Case 1: British Columbia shale Case 2: Barnett shale
H2S (ppmv) 26 750
CO2 (vol.%) 1.1 2.5
CH4 (vol.%) balance balance
T (°C) 31.8 32.2
P (MPa) 3.10 6.62
Flow (MMSCFD) 90 330
H2S removed (kg/h) 3.35 417
In both cases the gases should be treated to pipeline quality, i.e. 4 ppmv H2S and below 2 vol.% of CO2, according to Weiland and Hatcher.13
For Case 1, the only compound required to be removed is H2S, because CO2 is already below pipeline specifications. Therefore it is desirable to select a process with the lowest possible removal of CO2. In Case 2, both H2S and CO2 need to be treated to reach pipeline specifications. The Vitrisol® process is not able to remove CO2; therefore additional processing is required for case 2 to obtain on spec gas. For this purpose also an amine treating process is chosen.
For the acid gas injection of the strip gas the pressure was chosen to be 15 MPa (2176 psi) which was realized with a multi-stage compressor.
7.3 “Amine-treated” cases by PPS
7.3.1 Introduction to PPS
Alkanolamines have been widely used for more than 80 years in gas treating industry i.e. petrochemicals, refineries, natural gas processing. Recently, formulated amines that comprise of a promotor have been incorporated in gas treating and in large scale post-combustion CO2 capture. Acid gas treating industry mainly consists of processes where one or more gaseous components are transferred from the gas phase to the liquid phase followed by a chemical reaction. Due to the complexity of the absorption processes, modeling them requires very precise knowledge of reaction kinetics, mass transfer, thermodynamics and physical properties. In addition to the development of rigorous models that account for the aforesaid phenomena, it is also important to incorporate the correct description of vapor-liquid and liquid phase chemical equilibria including the speciation of the various components.
A steady-state rate based flowsheeting software for the simulation of acid-gas treating processes has been developed by Procede.14 The flowsheeting tool has models that can do the design, optimization and analysis of acid gas treating processes including both pre and post combustion CO2 capture, respectively. The process simulator consists of a user-friendly graphical user interface (GUI) and a powerful numerical solver that can handle rigorous simultaneous solutions of thermodynamics, kinetics, rate-based mass transfer equations (also known as rate-based model) and supports all unit
124
operations involved in gas treating such as absorbers, strippers, flash drums, heaters, pumps, compressors, mixers and splitters as well as work flow tools such as automatic water and solvent makeup calculators.
The Procede Process Simulator (PPS) has extensive, carefully evaluated databases of thermodynamic parameters, interaction coefficients, kinetics that have been optimized to accurately predict vapor liquid equilibrium (VLE), thermodynamic and physical properties and kinetically enhanced mass transfer (both approximate and rigorous) for amine- and mixture of amines based capture processes. PPS is able to describe complete gas treating processes involving complex flow schemes with multiple recycle loops. Both absorber and stripper can be modeled as rate-based columns. For optimal predictions of column performances, the program includes databases of various commercially available tray types and a large collection of both dumped and structured packings, several mass transfer and hydrodynamic correlations from open literature are implemented. PPS has capabilities where users can include detailed characterization of proprietary amines, mixtures of amines, mixtures of amine and physical solvents obtained from experiments used for the development of new gas treating processes.
7.3.2 Process description
The pipeline specifications of shale gas treating is to remove H2S to < 4 ppmv and CO2 down to 2.0 vol.%. As both acidic gases will be absorbed simultaneously usually excessive amounts of CO2 are removed. The major challenge is to treat the stripper gas which also has considerable amounts of CO2 which might be of inferior quality for a Claus plant.13 It is important to note that the H2S absorption process is usually mainly gas-phase resistance controlled and CO2 absorption is liquid-phase resistance controlled. Traditionally N-methyldiethanolamine (MDEA) has been the solvent of choice in terms of cost and effectiveness. The gas treating plant consisted a/o of an absorber, stripper, lean-rich heat exchanger, lean solvent cooler, reboiler, condenser and pumps. The absorber and stripper were modeled using the Higbie-penetration model for mass transfer. Structured packings were used in the absorber and the stripper and the absorber section is operated counter currently. Aqueous 50 wt.% MDEA solution was selected as the solvent for Case 2a and a stripper promoter was added. For Case 1, the solvent circulations rate was 20 m3/h and for Case 2 and Case 2a, it was 100 m3/h.
The dimensions of the absorber and stripper columns are presented in Table 7.2, structured packing Sulzer MellaPak Plus with a geometric surface of a = 250 m2/m3 was used. For the calculations of the mass transfer parameters the correlations of Brunazzi and Paglianti15 for the specific area, Bravo and Rocha16 for the liquid phase mass transfer coefficient and Bravo et al.17 for the gas phase mass transfer coefficient, respectively.
Table 7.2: Absorber and stripper dimensions.
Case 1 Case 2 Case 2a
British Columbia shale Barnett shale Barnett shale
Absorber
Packing Sulzer MellaPak Plus Sulzer MellaPak Plus Sulzer MellaPak Plus
Height (m) 12 25 25
Diameter (m) 1 1.35 1.35
Stripper
Packing Sulzer MellaPak Plus Sulzer MellaPak Plus Sulzer MellaPak Plus
Height (m) 3 10 10
125
7.3.3 PFD
Figure 7.1: PFD of the Amine process with regeneration.
7.3.4 Results
7.3.4.1 Case 1For the process conditions given in Table 7.1 and dimensions given in Table 7.2 the H2S-specification could be met. As a consequence of the H2S removal also CO2 was (partly) removed from 1.10 vol.% to 0.98 vol.%. The stream with a flow of 128 Nm3/h leaving the stripper (stream 6 in Figure 7.1) consists of 97 vol.% CO2 and 1.75 vol.% H2S and the stream that needs further processing. The H&M balances of Case 1 are given in Appendix 7.A.
Next, the utility consumptions were calculated. For pumps, an efficiency of 80 % was assumed. Heating can be performed with low pressure steam. Cooling is performed with either air (above 40 °C) or water (below 40 °C). For the heat integration lean-rich heat exchanger (E-101), a temperature approach of 5 °C was assumed. Though MDEA has a low volatility, some MDEA make-up is required. The electrical power consumption of the multi-stage compressor was added to the total electrical duty required for the process. By efficient compressor interstage cooling a part of the heat can be recovered and possibly reused in the amine process. The utility consumptions are given in Table 7.3.
Table 7.3: Utility consumptions of cases 1 and 2 of the Amine process.
Case 1 Case 2a
British Columbia shale Barnett shale
Electrical power (kW) 46.8 740
Heating duty (kW) 569a 5192a
Cooling duty (kW) 371b, 0c 2892b, 0c
MDEA (kg/h) 0.081 0.86
a: Reboiler duty
b: Cooling duty above 40 °C c: Cooling duty below 40 °C
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7.3.4.2 Case 2
For the process conditions given in Table 7.1 and dimensions given in Table 7.2, both the H2S- and CO2-specification could not be met. Therefore a stripping promoter has been added. In this particular case, i.e. Case 2a, H3PO4 with a concentration of 0.3 wt.% was used. It turned out that meeting the H2S-specification was the process limiting step, in the simulations an outlet concentration was attained of 3.92 ppmv. The CO2-specification of 2.0 vol.% was easily realized, the simulations gave an outlet concentration of 1.83 vol.%, so about 35 % more CO2 was removed than was demanded. The stream (gas) with a flow of 2175 Nm3/h leaving the stripper consists of 90.1 vol.% CO2 and 9.9 vol.% H2S, a gas outlet stream that certainly demands further processing. The H&M balances of Case 2a are given in Appendix 7.B. The utility consumptions are also given in Table 7.3.
It can be concluded that amine gas treating processes can be used for the upgrading of Cases 1 and 2 to meet the desired specifications. However, the obtained off-gases need further processing. In this study it was chosen to treat the off-gases with acid gas injection. The composition of the off-gases is such that treating with a Claus-operation is not feasible.
7.4 Vitrisol® process extended with regeneration of active component.
First an in-depth description of the Vitrisol® process will be provided which is necessary to understand process designs, the H&M balances and equipment designs for Cases 1 and 2, respectively.
7.4.1 Technology description
For medium and large scale H2S removal operations, as e.g. the described shale gas Cases in Table 7.1, the Vitrisol® process will be operated with a regeneration section to minimize operational costs (i.e. the consumption of copper). The Vitrisol® process with regeneration basically consists of three steps, i.e. the absorption section, the regeneration section, and the sulphur recovery section, respectively.
7.4.1.1 Parameters determining the process boundary conditions
The parameters that have a substantial impact on the process design of the Vitrisol® process with regeneration because of a/o issues with stable operation are:
1. The wt.% of copper sulphide (CuS) in the aqueous absorption liquid exiting the absorber, 2. The concentration of copper sulphate (CuSO4) in the aqueous absorption liquid exiting the
absorber.
From experiments performed with the previously mentioned Vitrisol® pilot absorber12 it was concluded that the rheological and hydraulic behaviour of the aqueous solution abruptly changed above concentrations of 3 wt.% to 5 wt.% of CuS. Furthermore, it was observed that the addition of Fe2+ to the absorption liquid (containing CuSO4 and H2SO4) did not affect the H2S removal efficiency in the absorber.
In the conceptual design exit concentrations of CuS of 0.25 wt.% and 1.0 wt.% were chosen for Cases 1 and 2, respectively. Flexibility and turndown options are introduced into the design with respect to operation by not designing the process near the maximum CuS concentration. E.g.:
1. Sudden spikes in H2S concentration will not affect the operability of the process significantly 2. A consistent, temporary increase in H2S concentration of the gas in time will not require
significant alteration of the installed hardware (a phenomena often encountered with e.g. shale gas)
127 From the work by Ter Maat et al.4,5 it can be concluded that the reaction rate of H2S with Cu2+ is instantaneous compared to the mass transfer rate, even at very low concentrations of Cu2+. Therefore, the removal of H2S is not a function of the concentration of Cu2+ in the absorption liquid. Moreover, from the work by Wermink and Versteeg6,9,10 regarding Fe2+ oxidation experiments performed in the presence of Cu2+ it followed that the presence of Cu2+ increased the conversion rate of Fe2+ with O2 to Fe3+ substantially compared to Fe2+ oxidation experiments performed without Cu2+. From the Fe2+ oxidation experiments performed in the presence of Cu2+ it was concluded that the presence of sulphuric acid (H2SO4) did not affect the conversion rate of Fe2+ to Fe3+, and that an increase in Fe2+ concentration increased the conversion rate of Fe2+ to Fe3+ more significantly than an increase in Cu2+ concentration.
From the experimental work it was concluded that the dissolution of CuS, in the presence of only Fe2+, Cu2+, H2SO4, and O2 (no Fe3+ initially), followed an initial conversion rate of CuS equal to the conversion rate of Fe2+ with O2 in the presence of Cu2+ up to roughly 80 % conversion. Therefore it was concluded that:
1. Up to a CuS conversion of roughly 80 % the oxidation of Fe2+ is the rate determining step in a batch reactor
2. It is desirable to increase the Fe2+ concentration instead of the Cu2+ concentration to obtain a higher conversion rate of Fe2+ in the oxidation step of the process
Concentrations of CuSO4, FeSO4 and H2SO4 of 0.05 M, 0.75 M and 0.1 M were chosen, respectively, for the absorption liquid exiting the absorber.
7.4.1.2 Absorption section
The absorption of H2S is completely limited by mass transfer in the gas phase and therefore 1st order with respect to the removal of H2S. This implies that most of the CuS solids are precipitating in the first section of a G/L contactor. Therefore in the selection of the absorber type a combination has been chosen of a packed bed and bubble column. Deep removal is realised in the packed bed section while the gas-liquid disengagement zone below the packed bed section is designed as a bubble column. This bubble column section is also acting as liquid storage vessel.
The bubble column section was designed based on:
1) The specific interfacial areas of the bubble column for Cases 1 and 2 were determined to be a = 190 m2/m3 and a = 250 m2/m3, respectively, see Oyevaar.18
2) The gas hold-ups of the bubble column for Cases 1 and 2 were determined to be ε = 0.35 and ε = 0.37, respectively, see Oyevaar.18
3) The gas phase mass transfer coefficient was calculated using the approach of Colombet,19 however, from this work it is not evident how the effect of process pressure must be taken into account as the gas phase diffusion coefficient is inversely proportional to the pressure. 4) The gas phase mass transfer coefficient was corrected for the pressure in two manners, i.e.
according to the film theory and Higbie’s penetration theory, respectively. According to film theory, the gas phase diffusion coefficient is inversely proportional with pressure, whereas the gas phase diffusion coefficient is inversely proportional with the root of pressure according to penetration theory.
5) The actual superficial velocity of the gas in the bubble column was set to vs,G = 0.15 m.s-1.
Although the absorption of H2S into the Vitrisol® liquid is irreversible and therefore no real preference exists for counter-current operation, this mode has been selected in the present designs.
128
1) As packing 1 inch Pall rings were selected. The effective interfacial area was calculated according to the correlations of Billet and Schultes.20
2) The gas phase mass transfer coefficient was calculated according to the correlations of Onda et al.21
3) The absorber operation was designed to reach H2S pipeline specification independent of the bubble column operation.
In Case 2, the pressure of the Vitrisol® liquid exiting the absorber is flashed to the pressure in the oxidizer reactor. The flash vessel is designed for a residence time of half an hour of the liquid. The flash vessel is stirred to maintain a homogeneous solid-liquid solution. In Case 1 the flash vessel is omitted, because CH4 losses are negligible.
Prior to entry in the oxidizer, heat is exchanged between the Vitrisol® liquid leaving the absorption section and regenerated Vitrisol® liquid entering this section.
7.4.1.3 Regeneration section
From the work by Wermink and Versteeg11 it was concluded that at all the temperatures investigated, i.e. temperatures ranging from 25 °C to 90 °C, a conversion of CuS of approximately 100 % could be obtained. Moreover, with temperature the rate of dissolution of CuS increased. For the components Fe3+ and H2SO4 both zero order dependencies were observed.
Based on CuS dissolution experiments at 90 °C, in the presence of CuS, FeSO4 and O2, it was concluded that the first 80 % conversion of CuS occurred at a rate identical to the initial conversion rate of Fe2+ with O2 in the presence of Cu2+. The latter 20 % could be fully converted, but required an additional residence time of approximately one hour. Furthermore, at a level of 80 % conversion of CuS sufficient amounts of Fe3+ were produced to convert the remaining 20 % of the CuS. This was independently confirmed experimentally in which the O2 supply was stopped at a conversion level of 80 %. Therefore it was concluded that continuous air flow was not required to fully dissolve the CuS and retain the Cu2+.
From oxidation experiments of Fe2+ with oxygen in the presence of Cu2+ it was concluded that some of the performed experiments were affected by mass transfer of O2 in the liquid.10 Therefore intrinsic kinetics could not be determined exactly. As the laboratory scale reactor was designed to have very high mass transfer rates, it can be concluded that for process scale types, with substantially lower kL and a values, the mass transfer of O2 to the liquid is the rate determining step of the CuS dissolution. Therefore the design of this unit is based on the absorption rate of O2.
The dissolution process of CuS will be carried out in three steps; i.e. an oxidation step, an extraction step and a candle filter operation, respectively.
In the oxidation step a continuous pressurized air flow is fed to a gas-liquid contactor with high-intensity stirring at a temperature of 90 °C and a pressure of 1 MPa. The heat generated by the compressor can be fully integrated in the Vitrisol® process. An excess amount of O2 is present to ensure a stable O2 partial pressure. For compressors, pumps and stirrers an efficiency of 80 % was assumed. Build-up of water can occur in the process, as the overall reaction is given by:
( )
Preferably, the water produced in the regeneration section is evaporated in the absorber. However, for Case 1 the liquid flow is relatively low compared to the gas flow, resulting in a negligible temperature increase of the gas and therefore no water evaporation. Consequently, for Case 1 the compressor was designed to provide a larger flow to both ensure a stable O2 pressure and evaporate sufficient amounts of H2O. For Case 2 the process has been designed to evaporate the excess water in the absorber, therefore the compressor is designed to only ensure a stable O2 pressure.
129 The basis of design of the oxidizer is the relation by Van ‘t Riet22 derived for kLa in stirred vessels with ionic aqueous solutions. A power to volume ratio of the stirrer of 2000 W/m3 is selected as design parameter, resulting in a kLa of 0.15 s-1 and 0.22 s-1 for Cases 1 and 2, respectively. The solubility of O2 in the Vitrisol® solution was determined with an adapted version of the model by Weisenberger and Schumpe.10,23
Subsequently the Vitrisol® liquid enters the extraction step, i.e. a liquid-liquid contactor with high-intensity stirring also at a temperature of 90 °C. In the liquid-liquid contactor an emulsion is created, elemental sulphur is dissolved and the dissolution of CuS proceeds in the absence of air. The amount of xylene entering the extractor is equal to 90 % of the maximum solubility of sulphur in xylene between a temperature of 50 °C and 90 °C to be able to dissolve an amount of sulphur equal to the sulphur present in the CuS formed in the absorber. The residence time in the liquid-liquid contactor of τ = 22.5 min was determined from experiments dissolving sulphur in xylene. A residence time of approximately 1 h for full conversion of CuS is not required, because the extraction step is operated continuously instead of batch-wise, resulting in increased conversion rates of CuS. A candle filter operation is placed subsequent to the extraction step to prevent possible slip of unconverted CuS to the crystallization section, because modes of operation could be possible that result in a conversion of CuS of below 100 %.
Subsequent to the candle filter operation the solids-free Vitrisol® liquid and xylene enter a settler to separate both phases. A residence time of 15 min is selected to separate both phases. The Vitrisol® liquid is returned to the absorber and xylene, partially saturated with dissolved sulphur, enters the crystallizer in the sulphur recovery section at a temperature of 90 °C. The dissolved xylene amount in the Vitrisol® liquid is determined from the solubility of p-xylene in water.24 Xylene, dissolved in the Vitrisol® liquid, is stripped in the absorber. Xylene concentrations in the gas leaving the absorber in Cases 1 and 2 are determined to be 2.0 ppmv and 16.6 ppmv, respectively.
7.4.1.4 Sulphur recovery section
In the crystallizer the xylene with the dissolved sulphur is decreased in temperature from 90 °C to 50 °C, resulting in the formation of sulphur crystals. Sulphur solubilities and dissolution times in various organic solvents, i.e. e.g. toluene, p-xylene, m-xylene and o-xylene, were experimentally determined and used in the design of the crystallizer. Afterwards, the xylene containing dissolved sulphur as well as crystallized sulphur is treated in a vacuum belt filter to separate the sulphur crystals from the liquid. The dry matter content of the sulphur crystals was assumed to be 70 wt.%. Xylene with dissolved sulphur is returned to the extraction step, sulphur cake is further processed in a sulphur melting operation. Sulphur cake is heated till the boiling point of xylene in the sulphur melting operation, i.e. a temperature of 140 °C. Xylene evaporates and is returned to the extraction step after condensation. 99+% liquid sulphur is stored in a storage tank. A make-up stream of xylene is fed to the extraction step because of xylene losses in the absorber.
It should be mentioned that the sulphur recovery section could be performed differently, when continuous removal of sulphur is not required, and the amount of sulphur to be removed is rather low (like e.g. Case 1). A saturation vessel could be included in the process design, i.e. a relatively large stirred vessel containing xylene with a sulphur storage capacity equal to a couple of weeks to months of operation. A possible mode of operation would be to remove (part of) the saturated xylene, and crystallizing the sulphur from the saturated xylene in vessel(s) in contact to the surroundings. Subsequently, the lean xylene can be returned to the Vitrisol® process, and the solid sulphur removed. In this type of process the use of continuous cooling and sulphur removal is not required.
130
7.4.1.5 CO2-absorber
For Case 2 the CO2 concentration is too high to meet the pipeline specifications (see Table 7.1). Therefore an additional CO2 removal technology must be applied. For the CO2 removal step also an amine based technology is chosen. It must be noted, however, that after the Vitrisol® process the H2S concentration is already 4 ppmv (or lower as will be discussed below) and therefore intrinsically no need exists to use a selective amine as MDEA. Moreover, the gas produced by the stripper is 99+% pure CO2 which can be re-used or directly vented to the environment depending on the local regulations. Therefore no compression of the off-gas for acid gas injection is required. Initially, an aqueous 50 wt.% MDEA (Case 2b) solvent has been selected for the removal of the CO2 to meet the specifications. Also, simulations by PPS have been carried out for an aqueous 30 wt.% MEA (Case 3) solution. For both Cases 2b and 3, the solvent circulation rate was 55 m3/h, respectively.
7.4.2 PFD
The PFD of the Vitrisol® process with regeneration for Case 1 is shown in Figure 7.2:
Figure 7.2: PFD of the Vitrisol® process with regeneration for Case 1.
131 Figure 7.3a: PFD of the Vitrisol® process with regeneration for Case 2 (H2S removal section).
The PFD of the CO2 removal process for Case 2 is shown in Figure 7.3b:
Figure 7.3b: PFD of the CO2 removal process for Case 2.
7.4.3 Results
The dimensions of the two-stage absorber (absorber column and bubble column), oxidation column and extraction column of the Vitrisol® process for Cases 1 and 2 are presented in Table 7.4, respectively.
132
Table 7.4: Two-stage absorber, oxidation and extraction column dimensions of the Vitrisol® process.
Case 1 Case 2
British Columbia shalea Barnett shaleb
Absorber column (C-101)
Packing 1 inch Pall rings 1 inch Pall rings
Height (m) 3 5.2 Diameter (m) 1.01 1.14 Bubble column (C-101) Height (m) 2.85 3.61 Diameter (m) 2.85 3.61 Oxidation column (R-101) Height (m) 1.51 7.26 Diameter (m) 0.50 2.42
Extraction column (V-101 and V-102)
Height (m) 2.75 8.68
Diameter (m) 0.92 2.89
a: Extraction column for Case 1 is V-101 b: Extraction column for Case 2 is V-102
The H&M balances of the Vitrisol® process for Cases 1 and 2 are given in Appendices 7.D and 7.E, respectively.
The utility consumptions are given in Table 7.5:
Table 7.5: Utility consumptions of Cases 1 and 2 of the Vitrisol® process.
a: Cooling duty above 40 °C
Because the Vitrisol® process can be operated at relatively mild conditions, heating can be performed with low pressure steam. Cooling can be performed with cooling water and/or air, because cooling duties above a temperature of 40 °C are required.
Xylene make-up is required because xylene losses arise from xylene exiting the process in the absorber. Xylene exiting the absorber is not considered to be a loss, because xylene is a component frequently encountered in natural gas. Xylene concentrations in the gas leaving the absorber in Cases 1 and 2 are determined to be 2.0 ppmv and 16.6 ppmv, respectively.
The utility consumptions for the removal of CO2 from the Vitrisol®-treated gas stream of Case 2 are given in Table 7.6b as Cases 2b and 3, respectively. From Tables 7.3 and 7.6b it can be observed that the utility consumption, especially the reboiler duty, has considerably reduced when H2S is removed upstream of the amine plant. Equipment specifications of the amine process are given in Table 7.6a.
Case 1 Case 2
British Columbia shale Barnett shale
Electrical power (kW) 11.4 784
Heating duty (kW) 18.9 520
Cooling duty (kW) 1.4a 178a
133 Table 7.6a: Absorber and stripper dimensions of the Amine process.
Case 2b: Barnett shale Case 3: Barnett shale
Absorber
Packing Sulzer MellaPak Plus Sulzer MellaPak Plus
Height (m) 25 17
Diameter (m) 1.35 1.38
Stripper
Packing Sulzer MellaPak Plus Sulzer MellaPak Plus
Height (m) 10 10
Diameter (m) 1.5 1.5
Table 7.6b: Utility consumptions of cases 2b and 3 of the Amine process. Case 2b: Barnett shale Case 3: Barnett shale
Electrical power (kW) 129 132
Heating duty (kW) 2140a 2060a
Cooling duty (kW) 607b, 0c 626b, 0c
Amine (kg/h) 0.79d 0.50e
a: Reboiler duty
b: Cooling duty above 40 °C c: Cooling duty below 40 °C d: MDEA
e: MEA
From the simulations no real preference for either two amines can be made based on operational costs. The dimensions of the MEA-based are slightly smaller compared to MDEA. In the comparison the data obtained for MDEA will be used.
7.5 Discussion
7.5.1 Comparison of amine treating solutions to Vitrisol®
The total utility consumptions for processing British Columbia shale to pipeline quality with standard amine treating solutions and Vitrisol® are given in Table 7.7.
Table 7.7: Utility consumptions of cases 1 and 2 of the Amine process.
Case 1 Case 1
Standard amine process Vitrisol® process
Electrical power (kW) 22.4 11.4
Heating duty (kW) 569a 18.9
Cooling duty (kW) 371b, 0c 1.4b, 0c
MDEA (kg/h) 0.081
Xylene (kg/h) 0.94
Xyleneabsorber gas (ppmv) 2.0
CO2,stripper gas (kg/h) 239 H2Sstripper gas (kg/h) 3.35 H2Sstripper gas (ppmv) 1.65 x 104 a: Reboiler duty
b: Cooling duty above 40°C c: Cooling duty below 40°C
134
From Table 7.7 it can be concluded that significant reductions with respect to utility consumptions can be achieved when Vitrisol® is used to remove H2S instead of a standard amine process. Furthermore, the H2S is oxidized in the Vitrisol® process, nullifying additional H2S removal operations downstream. E.g., the stripper gas in the amine process still contains 1.65 x 104 ppmv of H2S, as is shown in Table 7.7. Xylene make-up is not considered to be a loss, because xylene is a component frequently encountered in natural gas.
The total utility consumptions for processing Barnett shale to pipeline quality with standard amine treating solutions and a combination of Vitrisol® with a standard amine treating solution are given in Table 7.8.
Table 7.8: Utility consumptions of cases 1 and 2 of the Amine process.
Case 2a Case 2
Standard amine process Vitrisol® process + MDEA process
Electrical power (kW) 239 784 + 129 Heating duty (kW) 5191a 520 + 2140 Cooling duty (kW) 2892b, 0c 178b + 607b, 0c MDEA (kg/h) 0.86 0.79 Xylene (kg/h) 28.3 CO2,stripper gas (kg/h) 4889 3634 H2Sstripper gas (kg/h) 417 1.79 H2Sstripper gas (ppmv) 9.24 x 104 625 a: Reboiler duty
b: Cooling duty above 40°C c: Cooling duty below 40°C
From Table 7.8 it can be concluded that significant reductions with respect to utility consumptions can be achieved when Vitrisol® is used to remove H2S prior to CO2 removal with a standard amine process. Furthermore, though the H2S is oxidized in the Vitrisol® process, an additional H2S removal operation is required downstream to remove H2S from the amine stripper gas. As described in Table 7.8, the quantity of H2S to be removed from the stripper gas varies significantly between Barnett shale treated by a standard amine process for both H2S and CO2 removal and the combination of the Vitrisol® process for H2S removal and standard amine process for CO2 removal.
As the costs of energy (and cooling) are extremely location depended, no attempt was made to quantify the capital savings.
7.5.2 Enhanced H
2S removal of Barnett shale gas (case 2)
As explained before, the Barnett shale gas coming from the Vitrisol® process is required to be treated by a standard amine process to remove CO2. However, the feed gas for the amine absorber, containing 4 ppmv of H2S, will result in an increased H2S concentration coming from the amine stripper. If the H2S content of the feed gas for the amine absorber would be decreased below 4 ppmv, the possibility exists to produce a stripper gas which can be directly vented to the environment without additional H2S removal. The only requirement would be to provide additional packing in the Vitrisol® absorber to enhance the removal of H2S.
Figure 7.4 shows the amount of H2S in the gas leaving the stripper for varying of amounts of H2S in the feed gas.
135 Figure 7.4: H2S content of the gas exiting the stripper as a function of the H2S concentration in the feed gas.
Figure 7.5 shows the amount of H2S in the gas entering the amine absorber and the gas leaving the amine stripper for varying packing heights in the Vitrisol® absorber.
Figure 7.5: H2S content of gas entering the amine absorber (feed) and of gas exiting the amine stripper (out) as a function of packing height in the Vitrisol® absorber.
136
From Figures 7.4 and 7.5 it can be concluded that it is advantageous to remove more H2S upstream with the Vitrisol® process than is required for the gas specifications as this enables an easy handling downstream of the amine stripper gas, e.g. venting CO2 to the environment. From Figure 7.5 it can be concluded that a packing height of 3.4 m is required to meet the 4 ppmv specifications, however, this will lead to a H2S concentration in the CO2 off-gas from the amine stripper of 625 ppmv. To arrive at a concentration of H2S of about 1 ppmv additionally about 4 m packing is required (about 4 m3).
7.6 Conclusions
Procede developed over the recent years a solvent that is 100% selective for H2S removal from industrial gases in the presence of CO2, called Vitrisol®. Vitrisol® is able to remove in one stage more than 99.9+% of the H2S present in the gas phase and has the typical characteristics of very selective H2S scavengers. However, a major difference of Vitrisol® compared to the traditional scavengers is that Vitrisol® can be completely regenerated resulting in a solvent with fully restored activity and crystalline sulphur.
For the absorption process required for the removal of H2S, there is no real process conditional constraint and the operating pressures and temperatures can vary at least between 0.1-10 MPa and 283-363 K, respectively. The Vitrisol® regeneration process takes place at temperatures below 373 K. The process pressure can vary from atmospheric up to 0.5-1 MPa. The Vitrisol® process can be described with the following overall reaction equation:
( )
In the Vitrisol® process no need of vast amounts of energy are required for the regeneration of the solvent, therefore this process has an extremely low energy footprint.
In the present contribution the performance of Vitrisol® is demonstrated for two applications in shale gas production as previously described by Weiland and Hatcher, i.e. British Columbia gas (Case 1) and Barnett gas (Case 2), respectively. The Vitrisol® process is also compared to a standard amine treating process designed for selective H2S removal. The remainder toxic, acid gas produced in the stripper of the amine treating process is compressed for acid gas injection.
From the results it can be concluded that significant reductions in operational costs can be achieved by using the Vitrisol® process as depicted in the energy consumptions of the overall process. E.g. in Case 1 the total energy consumption of the Vitrisol® process is 32 kW, whereas the total energy consumption of the standard amine treating process is 987 kW. In Case 2 the total energy consumption of the Vitrisol® process in combination with a standard amine treating process is 4358 kW, whereas the total energy consumption of the standard amine treating process is 8824 kW. Moreover, no additional treatment of the off-gas stream is required for the process with Vitrisol® upstream of the amine treating unit, as the H2S is directly converted to crystalline sulphur. Depending on the local governmental regulation the off-gas could be directly vented to the environment. This study also illustrates clearly that it is advantageous to first remove H2S from a gas stream containing both H2S and CO2 prior to CO2 removal to reduce operational costs. As the costs of energy (and cooling) are extremely location depended, no attempt was made to quantify the capital savings.
137
7.7 Nomenclature
a interfacial area [m2/m3] D diameter [m] e hold-up [-] H height [m]k mass transfer coefficient [m.s-1]
P pressure [Pa, psi]
T temperature [°C, K]
vs superficial velocity [m.s-1]
Subscripts and superscripts
G gas
L liquid
7.8 References
[1] Mcintyre, G. and Lyddon, L. Claus sulphur recovery options. Pet. Technol. Q. Spring 1997, 57─61. [2] Perry, D., Fedich R.B. and Parks, L.E. Better acid gas enrichment. Sulphur 2010, 326, 38─42. [3] Versteeg, G.F. and Ter Maat, H. Method and system for selective removal of contamination from
gas flows. WO patent 1998055209 A1, assigned to Procede Twente B.V., priority date June 2, 1997.
[4] Ter Maat, H., Hogendoorn, J.A. and Versteeg, G.F. The removal of hydrogen sulfide from gas streams using an aqueous metal sulfate absorbent. Part I. The absorption of hydrogen sulfide in metal sulfate solutions. Sep. Purif. Technol. 2005, 43 (3), 183─197.
[5] Ter Maat, H., Al-Tarazi, M., Hogendoorn, J.A., Niederer, J.P.M. and Versteeg, G.F. Theoretical and experimental study of the absorption rate of H2S in CuSO4 solutions. The effect of enhancement of mass transfer by a precipitation reaction. Chem. Eng. Res. Des. 2007, 85 (1), 100─108.
[6] Wermink, W.N. and Versteeg, G.F. The oxidation of Fe(II) in acidic sulphate solutions with air at elevated pressures. Part 1. Kinetics above 1 M H2SO4. Ind. Eng. Chem. Res. 2017, 56 (14), 3775– 3788.
[7] Peacey, J., Guo, X.-J. and Robles, E. Copper hydrometallurgy – current status, preliminary economics, future direction and positioning versus smelting. Trans. Nonferrous Met. Soc. China
2004, 14 (3), 560─568.
[8] Dutrizac, J.E. and MacDonald, R.J.C. The kinetics of dissolution of covellite in acidified ferric sulphate solutions. Can. Metall. Q. 1974, 13 (3), 423─433.
[9] Wermink, W.N. and Versteeg, G.F. The oxidation of Fe(II) in acidic sulphate solutions with air at elevated pressures. Part 2. Influence of H2SO4 and Fe(III). Ind. Eng. Chem. Res. 2017, 56 (14), 3789–3796.
[10] Wermink, W.N., Spinu, D. and Versteeg, G.F. The oxidation of Fe(II) with Cu(II) in acidic sulphate solutions with air at elevated pressures. Chem. Eng. Commun. 2018, accepted.
[11] Wermink, W.N. and Versteeg, G.F. The dissolution of CuS with Fe(III) in acidic sulphate solutions.
Ind. Eng. Chem. Res. 2018, submitted.
[12] Ter Maat, H., Versteeg, G.F. and Vergunst, F. Sour Oil & Gas Advanced Technology 2012, Proceedings of the 8th International Conference, Abu Dhabi, UAE, March 27–28, 2012.
[13] Weiland, R.H. and Hatcher, N.A. Overcome challenges in treating shale gas. Hydrocarb. Process.
2012, 91 (1), 45–48.
[14] Van Elk, E.P., Arendsen, A.R.J. and Versteeg, G.F. A new flowsheeting tool for flue gas treating.
Energy Procedia 2009, 1 (1), 1481–1488.
[15] Brunazzi, E. and Paglianti, A. Liquid-film mass-transfer coefficient in a column equipped with structured packings. Ind. Eng. Chem. Res. 1997, 36 (9), 3792–3799.
138
[16] Bravo, J.L. and Rocha, J.A. A comprehensive model for the performance of columns containing structured packings. IChemE Symp. S. 1992, 128, A439-A457.
[17] Bravo, J.L., Rocha, J.A. and Fair, J.R. Mass transfer in gauze packings. Hydrocarb. Process. 1985, 64
(1), 91–95.
[18] Oyevaar, M.H. and Westerterp, K.R. Interfacial areas and gas hold-ups in gas-liquid contactors at elevated pressures from 0.1 to 8.0 MPa. Chem. Eng. Sci. 1991, 46 (5/6), 1217–1231.
[19] Colombet, D., Legendre, D., Cockx, A. and Guiraud, P. Mass or heat transfer inside a spherical gas bubble at low to moderate Reynolds number. Int. J. Heat Mass Tran. 2013, 67, 1096–1105. [20] Billet, R. and Schultes, M. Prediction of mass transfer columns with dumped and arranged
packings. Chem. Eng. Res. Des. 1999, 77 (6), 498–504.
[21] Onda, K., Takeuchi, H. and Okumoto, Y. Mass transfer coefficients between gas and liquid phases in packed columns. J. Chem. Eng. Jpn. 1968, 1 (1), 56–62.
[22] Van ’t Riet, K. Review of measuring methods and results in nonviscous gas-liquid mass transfer in stirred vessels. Ind. Eng. Chem. Process Des. Dev. 1979, 18 (3), 357–364.
[23] Weisenberger, S. and Schumpe, A. Estimation of gas solubilities in salt solutions at temperatures from 273 K to 363 K. AIChE J. 1996, 42 (1), 298─300.
[24] Knauss, K.G. and Copenhaver, S.A. The solubility of p-xylene in water as a function of temperature and pressure and calculated thermodynamic quantities. Geochim. Cosmochim. Ac.
139
Appendix 7.A: H&M balance of case 1 (British Columbia shale) of the Amine Process
140
Appendix 7.B: H&M balance of case 2a (Barnett shale) of the Amine Process with
stripper promoter
141
Appendix 7.C: H&M balance of case 3 (Barnett shale) of the Amine Process (MEA)
142
Appendix 7.D: H&M balance of case 1 (British Columbia shale) of the Vitrisol®
process
St re am n u m b e r 1 2 3 4 5 6 7 8 9 10 11 12 N G in le t N G o u tl e t H P Cu S ri ch V itr is o l H P le an V itr is o l LP H T Cu S ri ch V itr is o l A ir Co mp re ss e d a ir V itr is o l Co mp re ss e d a ir W ate r A ir V itr is o l & x yl e n e Fl o w [ton /h ] 73 .3 73 .3 3. 76 3. 76 3. 76 0. 04 7 0. 04 7 3. 80 0. 04 9 0 0. 04 9 3. 9 Fl o w [a ct m3 /h ] 3440 3440 3. 62 3. 62 3. 79 39 5. 0 3. 79 5. 4 0 54 3. 9 Pr e ssu re [b ar a] 31 31 31 31 10 1. 01 10 10 10 1. 01 10 Te m p e ra tu re [° C] 31 .8 32 .2 31 .8 37 .5 90 15 90 90 90 90 90 Ph ase G G L+ S L L+ S G G L+ S G L G L C o n c.so li d s [w t. % ] 0 0 0. 25 0 0. 25 0 0 0. 11 0 0 0 [H2 S] [p p m] 26 4 n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . H2 O kg /h r 168 168 3308 3265 3308 0. 5 0. 5 3308 2 2 3308 H2 SO 4 kg /h r 45 36 45 36 36 C u SO 4 kg /h r 29 45 29 42 45 Fe SO 4 kg /h r 413 413 413 407 413 Fe 2 (S O4 )3 kg /h r 8 H2 S kg /h r 4. 0 0. 61 C u S kg /h r 9. 4 9. 4 1. 8 S kg /h r 2. 5 5. 1 CH 4 kg /h r 70957 70955 1. 62 1. 62 CO 2 kg /h r 2165 2165 O2 kg /h r 11 11 9 9 N2 kg /h r 35 35 35 35 X yl e n e kg /h r 0. 9 0. 9 55143 St re am n u m b e r 13 14 15 16 17 18 19 20 21 22 23 24 V itr is o l & x yl e n e V itr is o l X yl e n e X yl e n e Su lp h u r ca ke X yl e n e X yl e n e ma ke -u p X yl e n e V itr is o l ma ke -u p X yl e n e Li q u id s u lp h u r X yl e n e Fl o w [ton /h ] 3. 9 3. 8 0. 05 9 0. 05 9 4. 44 E-03 0. 05 4 9. 35 E-04 0. 05 7 0 0. 00 13 3. 15 E-03 1. 28 E-03 Fl o w [a ct m3 /h ] 3. 9 3. 8 0. 07 0. 07 3. 09 E-03 0. 06 1. 08 E-03 0. 06 8 0 0. 41 1. 75 E-03 1. 71 E-03 Pr e ssu re [b ar a] 1 31 1 1 1 1 10 10 1 1 1 Te m p e ra tu re [° C] 90 90 90 50 50 50 15 90 140 140 140 Ph ase L L L L+ S S+ L L L L L G L L C o n c.so li d s [w t. % ] 0 0 0 5. 5 70 0 0 0 0 0 0 0 [H 2 S] [p p m] n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . H2 O kg /h r 3308 3308 H2 SO 4 kg /h r 36 36 C u SO 4 kg /h r 45 45 Fe SO 4 kg /h r 413 413 Fe 2 (S O4 )3 kg /h r H2 S kg /h r C u S kg /h r S kg /h r 5. 1 5. 1 5. 1 3. 2 2. 0 2. 0 3. 2 CH 4 kg /h r CO 2 kg /h r O2 kg /h r N2 kg /h r X yl e n e kg /h r 55 0. 94 54 54 1. 3 52 0. 9 55 1. 3 1. 3
144
Appendix 7.E: H&M balance of case 2 (Barnett shale) of the Vitrisol® process
St re am n u m b e r 1 2 3 4 5 6 7 8 9 10 11 12 13 N G in le t N G o u tl e t H P Cu S ri ch V itr is o l H P le an V itr is o l Fl as h g as LP Cu S ri ch V itr is o l LP H T Cu S ri ch V itr is o l A ir Co mp re ss e d a ir V itr is o l Co mp re ss e d a ir W ate r A ir Fl o w [ton /h ] 27 5. 4 27 4. 8 11 7. 0 11 6. 6 0. 10 8 11 6. 9 11 6. 9 4. 28 4. 28 11 7. 0 4. 24 0 4. 24 Fl o w [a ct m3 /h ] 5534 5524 11 0. 8 11 0. 8 13 .4 11 0. 7 11 3. 4 3513 44 8. 6 11 3. 5 45 7. 8 0 4518 Pr e ssu re [b ar a] 66 .2 66 .2 66 .2 66 .2 10 10 10 1. 01 10 10 10 1. 01 Te m p e ra tu re [° C] 32 .2 35 .4 32 .3 38 .0 32 .1 32 .1 90 15 90 90 90 90 Ph ase G G L+ S L G L+ S L+ S G G L+ S G L G C o n c.so li d s [w t. % ] 0 0 1 0 0 1 1 0 0 0. 46 0 0 [H2 S] [p p m] 749 4 n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . H2 O kg /h r 427 427 99998 100074 0. 75 99997 99997 45 .0 45 .0 100072 191 191 H2 SO 4 kg /h r 2287 1086 2287 2287 1086 C u SO 4 kg /h r 884 2838 884 884 2462 Fe SO 4 kg /h r 12619 12619 12619 12619 11903 Fe 2 (S O4 )3 kg /h r 942 H2 S kg /h r 419 2. 24 C u S kg /h r 1170 1170 1170 225 S kg /h r 317 CH 4 kg /h r 256470 256346 124 10 7. 1 16 .7 16 .7 16 .7 16 .7 CO 2 kg /h r 18050 18050 O2 kg /h r 979 979 783 783 N2 kg /h r 3195 3195 3195 3195 X yl e n e kg /h r 28 .3 28 .3
145 St re am n u m b e r 14 15 16 17 18 19 20 21 22 23 24 25 26 V itr is o l & x yl e n e V itr is o l & x yl e n e V itr is o l X yl e n e X yl e n e Su lp h u r ca ke X yl e n e X yl e n e ma ke -u p X yl e n e V itr is o l ma ke -u p X yl e n e Li q u id s u lp h u r X yl e n e Fl o w [ton /h ] 12 4. 0 12 4. 0 11 6. 6 7. 32 7. 32 0. 55 6. 77 0. 02 8 6. 95 0 0. 16 0. 39 0. 16 Fl o w [a ct m3 /h ] 12 1. 9 12 1. 9 11 3. 6 8. 35 8. 19 0. 38 7. 80 0. 03 3 8. 39 0 51 .0 0. 22 0. 21 Pr e ssu re [b ar a] 10 1 66 .2 1 1 1 1 10 10 1 1 1 Te m p e ra tu re [° C] 90 90 90 90 50 50 50 15 90 140 140 140 Ph ase L L L L L+ S S+ L L L L L G L L C o n c.so li d s [w t. % ] 0 0 0 0 5. 5 70 0 0 0 0 0 0 0 [H 2 S] [p p m] n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . n .a . H2 O kg /h r 100072 100072 100072 H2 SO 4 kg /h r 1086 1086 1086 C u SO 4 kg /h r 2838 2838 2838 Fe SO 4 kg /h r 12619 12619 12619 Fe 2 (S O4 )3 kg /h r H2 S kg /h r C u S kg /h r S kg /h r 639 639 639 639 393 246 246 393 CH 4 kg /h r CO 2 kg /h r O2 kg /h r N2 kg /h r X yl e n e kg /h r 6708 6708 28 .3 6680 6680 160 6520 28 .3 6708 160 160
