investigation of process condensate treatment of a methane steam …€¦ · · 2009-06-02steam...
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KET050, Feasibility Studies on Industrial Plants Lund, 2009-05-29 Department of Chemical Engineering Lund University
Investigation of process condensate
treatment of a methane steam reforming plant
Tutors:
Almqvist, Robert Odenbrand, Ingemar, Dept. Chem. Eng., Lund University
Arkell, Anders Andersen, Kim H., Haldor Topsoe A/S
Fogel, Sebastian
Hansson, Rasmus
Söderström, Olle
i
Summary Steam reforming of methane need large amounts of steam, the non reacted steam is condensed in
the process. This condensate can be heated to steam again and be sold, for example for power
production. One large problem with the condensate is that it contains impurities formed in the
reformer and the shift converter, like ammonia, methanol and carbon dioxide which can damage
downstream equipment. To avoid this problem the condensate has to be treated.
Several different alternatives for cleaning are available, such as different biological ones, stripping
and ion-exchange. Due to the temperature, pressure and long residence time the biological
treatment was rejected. Steam stripping and ion-exchange are the ones chosen for further
investigations. This report also investigates how much of the contaminants that can be produced in
the shift reactor, simulated with the program MATLAB. The stripper configurations have been
simulated using the flow sheeting program Aspen Plus.
The condensate cleaning is done by three different configurations of strippers with a possibility of
adding an ion-exchanger as an extra treatment unit after the strippers. The different configurations
are steam stripping in a single column, a stripper with a reboiler and two steam stripping columns
where pH is changed between the two strippers by addition of caustic. The reason behind changing
the pH is that ammonia is at equilibrium with ammonium in a water solution. By raising the pH the
equilibrium is shifted towards ammonia which then can be stripped off.
If no ammonia can be accepted in the condensate an ion-exchanger unit is an alternative as a
complement to the stripper. Due to the very effective cleaning in the strippers no ion-exchanger is
included in the investigated configurations. The stripper with a reboiler configuration was rejected
because it was less efficient than the steam stripper for the same amount of energy used and that
the steam is already available at the plant. The design factors of interest are the number of trays in
the columns and the amount of steam needed.
Our simulations show that it is possible to reach a very efficient purification of the condensate, above
99.9 %. Using steam stripping can give a condensate with very low levels of the impurities
investigated, levels as low as 10-34 ppmw is reached for carbon dioxide. The total cost for this
configuration is 21 million USD over a ten year period. Adding caustic is very effective to drive of
ammonia, levels of 10-5 ppmw is reached. The adding of caustic hampers the removal of carbon
dioxide, still very low levels of the carbon dioxide is reached, 0.087 ppmw, at a total cost of 24 million
USD over a ten year period. Methanol is mainly dependent on the amount steam added and is
proven more difficult to remove than the other two, levels of around 1 ppmw is reached for the
configurations above. A summary of the result is shown in the tables below.
The choice of configuration depends on the purification requirements for the different contaminants
and the price of the steam used.
ii
Table i Total cost for different purification efficiencies and stripper configurations. The calculations are based on ten years depreciation period with 15 % internal rate of return.
Optimal
single column
Optimal
two columns
99 %
alt. 1
99 %
alt. 2
Configuration 1 column
34 trays
2 columns
9+17 trays
caustic added
1 column
17 trays
1 column
25 trays
Conc. in stripper effluent
(CO2/NH3/MeOH) [ppmw]
(6.8·10-34/
0.14/0.71)
(0.087/
3.4·10-5/1.2)
(7.9·10-15/
9.6/49)
(7.2·10-24/
5.4/32)
Total cost steam included [USD] 20 666 000 23 526 000 19 720 000 17 890 000
Total cost steam excluded [USD] 2 111 000 5 495 000 1 160 000 1 596 000
Table i cont.
99.9 %
alt. 1
99.9 %
alt. 2
99.9 %
alt. 3
Configuration 1 column 25 trays 1 column 34 trays 2 columns 9+17 trays
caustic added
Conc. in stripper effluent
(CO2/NH3/MeOH) [ppmw] (2.5·10-25/1.1/5.9) (5·10-34/1.1/7.3) (3.8/1.1·10-4/3.4)
Total cost steam included [USD] 20 150 000 18 410 000 21 490 000
Total cost steam excluded [USD] 1 596 000 2 111 000 6 465 000
An interesting configuration is two columns of each ten ideal trays, instead of the 5 + 10 investigated
in this report, and caustic added to the second column. This should allow very low levels of both
carbon dioxide and ammonia to be reached. The investment cost is not that much higher compared
to a single 20 ideal trays column. This configuration was not tested in this report. If further test are to
be done we recommend testing this configuration.
iii
Contents Summary .................................................................................................................................................. i
1. Introduction ..................................................................................................................................... 1
Literature study part ............................................................................................................................... 1
2. The process ...................................................................................................................................... 1
2.1. Feed gas ................................................................................................................................... 2
2.2. Feed purification ..................................................................................................................... 3
2.3. Methane reformer ................................................................................................................... 4
2.3.1. The Haldor Topsoe A/S solutions .................................................................................... 5
2.4. Water gas shift reaction .......................................................................................................... 6
2.5. Off-gas purification .................................................................................................................. 7
2.5.1. Pressure swing adsorption .............................................................................................. 8
3. Kinetics and theory of formation of side products ......................................................................... 9
3.1. Shift reaction with the surface intermediate HCOOH ............................................................. 9
3.2. Ammonia ............................................................................................................................... 10
3.3. Methanol ............................................................................................................................... 11
3.4. Methylamines ........................................................................................................................ 13
4. Influence of contaminants............................................................................................................. 14
4.1. Ammonia ............................................................................................................................... 14
4.2. Carbon dioxide ...................................................................................................................... 14
4.3. Carboxylic acids ..................................................................................................................... 15
4.4. Oxygen ................................................................................................................................... 15
5. Condensate treatment .................................................................................................................. 15
5.1. Biological nitrogen treatment ............................................................................................... 15
5.1.1. Nitrification .................................................................................................................... 15
5.1.2. Denitrification ................................................................................................................ 16
5.1.3. Other treatment processes ........................................................................................... 16
5.2. Removal of methanol with a membrane bioreactor ............................................................. 16
5.3. Stripping ................................................................................................................................ 16
5.4. Deaeration ............................................................................................................................. 19
5.4.1. Pressure Deaerators ...................................................................................................... 19
5.5. Ion exchange ......................................................................................................................... 20
5.5.1. Mordenite ...................................................................................................................... 21
5.5.2. Clinoptilolite .................................................................................................................. 21
iv
6. Discussion of the literature part .................................................................................................... 21
Simulation part ...................................................................................................................................... 22
7. Result and discussion .................................................................................................................... 22
7.1. Shift simulation ...................................................................................................................... 22
7.2. Stripping ................................................................................................................................ 23
7.2.1. Steam stripper ............................................................................................................... 25
7.2.2. Stripper with reboiler .................................................................................................... 28
7.2.3. Steam stripper with pH adjusting .................................................................................. 33
7.3. Ion exchanger ........................................................................................................................ 38
7.4. Verification of simulation results .......................................................................................... 39
7.5. Discussion .............................................................................................................................. 42
8. Economy ........................................................................................................................................ 43
9. Discussion ...................................................................................................................................... 44
10. Conclusion ................................................................................................................................. 45
11. References ................................................................................................................................. 46
Appendix .................................................................................................................................................. a
A. Example of Aspen results ................................................................................................................ a
B. Results .............................................................................................................................................. c
Steam stripper .................................................................................................................................. c
Stripper with reboiler ...................................................................................................................... e
Steam stripper with pH adjustment ................................................................................................. i
C. Economy calculations ..................................................................................................................... m
Material for the ion exchangers ..................................................................................................... m
Material for the strippers ............................................................................................................... m
Total cost ......................................................................................................................................... o
D. Calculation method in Aspen .......................................................................................................... q
1
1. Introduction Hydrogen is an important raw material for the chemical industry. One of the most important uses is
in the synthesis of ammonia. Other uses are in the refinery industry for hydrogenation of heavier oil
fractions. Another very important process where hydrogen is a key component is the production of
synthesis gas. [1]
Most of the hydrogen produced today uses hydrocarbons as raw materials. There are several ways of
converting the hydrocarbons, for example gasification and reforming. These main production
alternatives can also be divided into subgroups. An alternative, but expensive, production route is via
electrolysis of water. [2]
When producing hydrogen by steam reforming water condensate is produced. This condensate can
be sent to a boiler and exported to other processes in the plant as steam. In some processes, for
example turbines, it is essential that the level of contaminants in the steam is kept to a minimum.
The objective of the project is twofold. At first, a study of the formation of contaminants and possible
reaction kinetics for their formation in the shift reactor is performed. Secondly a study of possible
condensate treatment steps required to meet the steam specifications from the consumers is done.
The aim is to be able to describe the amount of contaminants produced and propose efficient ways
of removing them from the condensate. To limit the investigations the focus lies on what is assumed
to be the three impurities of greatest importance. These are ammonia, carbon dioxide and methanol.
Literature study part 2. The process Hydrogen can be produced by steam reforming from several different raw materials which result in a
lot of variations in the process, depending on the size of the hydrocarbons that are to be reformed.
But even if the raw material is the same, there are a lot of variations in the process depending on the
intended use of the produced hydrogen, the size of the plant, investment cost, the age of the plant
and so on.
The general steps are that the feed gas first goes through a purification step before it enters the
steam reformer in which the hydrocarbons reacts with steam to form carbon monoxide and
hydrogen. After that the gas passes to the water shift reactor where carbon monoxide reacts with
water to form carbon dioxide and hydrogen. The last step or steps are the off-gas purification. [3, 4]
A typical plant design is shown in figure 1. [5]
2
Figure 1 Hydrogen production by steam reforming, gas purification by pressure swing adsorption a) Desulfurization; b) Feed preheater/superheater; c) Reformer; d) Waste-heat boiler; e) CO shift reactor (HT shift); f) Cooling of raw gas; g) Pressure swing adsorption; h) Off-gas buffer for fuel; i) Convective zone with steam production, steam superheating, and air preheating. [5]
2.1. Feed gas Most feed stocks contain higher hydrocarbons. Higher hydrocarbons are unwanted in the main
reforming units because of higher energy costs due to the need for a higher steam to carbon ratio.
The hydrocarbons are converted to carbon oxides, hydrogen and methane in a pre-reformer. [6]
The feed gas is often natural gas. Natural gas has different composition depending on the source. A
typical composition is found in table 1 and limits of the elements that can be found in natural gas are
in table 2.
Table 1 Typical composition of natural gas. [7]
Natural gas composition Mole fraction Mass fraction
Methane (CH4) 0.9229 0.8437
Ethane (C2H6) 0.0360 0.0623
Propane (C3H8) 0.0080 0.0206
Butane (C4H10) 0.0029 0.0099
Pentane (C5H12) 0.0013 0.0053
Hexane (C6H14) 0.0008 0.0039
Carbon dioxide (CO2) 0.0100 0.0252
Nitrogen (N2) 0.0180 0.0289
Water (H2O) 0.0001 0.0001
Table 2 Limits for each component in natural gas. [7]
Component Mole fraction
Main components
Methane (CH4) 0.70
Nitrogen (N2) 0.20
3
Carbon dioxide (CO2) 0.20
Ethane (C2H6) 0.10
Propane (C3H8) 0.035
Butane (C4H10) 0.015
Pentane (C5H12) 0.005
Hexane (C6H14) 0.001
Heptane (C7H16) 0.0005
Octane and above (C8+) 0.0005
Hydrogen (H2) 0.10
Carbon monoxide (CO) 0.03
Helium (He) 0.005
Water (H2O) 0.00015
Minor and trace components
Ethylene (C2H4) 0.001
Benzene (C6H6) 0.0005
Toluene (C6H5CH3) 0.0002
Argon (Ar) 0.0002
Hydrogen sulphide (H2S) 0.0002
Oxygen (O2) 0.0002
Total unspecified components 0.0001
2.2. Feed purification The feed almost always contains impurities that could either act as poisons for the catalyst or react
with the reactor walls. It is therefore necessary to remove them before letting the feed gas into the
reactor.
The table below list some of the impurities that the feed gas can contain and hence have to be
removed before the gas enters the reformer.
Table 3 Feed gas impurities. [8]
Name Form
Sulfur H2S
COS
Mercaptans
Organic sulfides
Organic disulfides
Thiophenes
Chlorine HCl
Organic chlorides
Arsenic Arsine
Mercury Organo-mercury
Mercury
The most common of these are organic and inorganic sulphur, followed by chlorine compounds and
the most uncommon are metal compounds.
4
Mercury which is a small aerosol present in the natural gas is removed directly when the gas is taken
from the ground by a centrifugal method.
Sulphur, which is the biggest problem, is removed from the feed gas in two steps in order to get
below the level of two ppmv, which is the limit in order to protect the steam methane reforming
(SMR) catalyst. The first step in the desulphurization process is that the gas, with a temperature of
360 °C to 400 °C, enters a catalyst consisting of Co-Mo or Zn oxides. This hydrogenates the organic
sulphur compounds into H2S. After this the H2S has to be removed, which can be made in two
different ways. The first way is by absorption. The H2S absorbs in monoethanolamine (MEA),
methyldiethanolamine (MDEA) or purisol. This is the most common way to eliminate the sulphur
from the feed gas. One problem is that these processes cannot be used at high temperatures. In that
case another way is better. It is quite new and uses adsorption on activated carbon or zinc oxide. The
temperature in this case can be between 350 °C and 450 °C, which means that no extra cooling and
heating is needed. The reaction for H2S on zinc oxide is as follows:
𝑍𝑛𝑂 + 𝐻2𝑆 → 𝑍𝑛𝑆 + 𝐻2𝑂 (2.2.-1)
The zinc oxide can, after saturation, be regenerated with hot air at a temperature of 700 °C. The
sulphur is then converted to SO2, which is removed from the air with absorption.
Chlorine, usually in the form of HCl but also in organic compounds, is removed by a similar process as
the sulphur. The organic compounds of chlorine are reacting on the same catalyst as organic sulphur
compounds to produce HCl. The hydrogen chloride is then absorbed. [3, 8, 9]
2.3. Methane reformer After purification the feed gas enters the reformer at a pressure of 25 to 35 atm
and the temperature is raised to between 780 °C and 900 °C. The catalyst
consists of alkali-promoted nickel and facilitates the endothermic reformation
reaction:
𝐶𝐻4 + 𝐻2𝑂 → 𝐶𝑂 + 3𝐻2 (∆H=227 kJ/mol) (2.2-2)
Because of the fact that the reaction is highly endothermic, the reactor has to be
heated by combustion of both recirculated waste gas from the downstream
purification as well as fresh natural gas. [3]
The choice of material in the reformer tubes is essential as it must withstand stress due to the high
temperature and temperature gradient. One possible material is a 35/25 Ni/Cr alloy modified with
niobium and micro alloyed with trace elements such as titanium and zirconium. [11]
In table 4 are some of the possible unwanted reactions of interest in this report listed.
Figure 2 Reformation reaction. [10]
5
Table 4 Possible and unwanted reactions in the steam reformer. [12]
Product Reaction stoichiometry Required H2/CO ratio
Methanol 2 H2 + CO → CH3OH 2.0
Acetic acid CH3OH + CO → CH3COOH 0
Methyl formiate CH3OH + CO → CH3OOCH 0
Methyl amines [13] NH3 + CH3OH ⇄ CH3NH2 + H2O
(CH3)2NH+CH3OH ⇄ (CH3)3N+H2O -
Hydrogen cyanide [14] CxH2x+2 + xNH3 → xHCN + 2(2x+1)H2 -
Elementary carbon could be formed and then deposited on the catalyst, this
can be avoided if a high steam to methane ratio is used or if the catalyst is
coated with CeO2. It is important to avoid carbon deposition because it will
deactivate the catalyst. [10]
2.3.1. The Haldor Topsoe A/S solutions
The capacity of Topsoe´s steam reformers range from a few 100 Nm3/h to
over 200 000 Nm3/h of pure hydrogen. [15, 16]
2.3.1.1. Tubular reforming
The Topsoe reformer is a side-fired reformer where the tubes are placed in a
single row inside the furnace chamber. [17] The burners heat the furnace
walls to a couple of thousand degrees, figure 3. The heat is then transferred
from the hot walls to the tubes via radiation. Radiation stands for almost all
heat transfer, about 90 %. [18] This reformer is designed for a high heat flux
and is thus called High Flux Reformer (HFR). The HFR can withstand tube
temperatures up to 1050 °C. This is the reformer with the highest capacity,
over 200 000 Nm3/h of pure hydrogen. [16]
A variant of tubular reformer is the Topsoe Bayonet Reformer (TBR). The
TBR consists of two tubes, the reformer tube which is closed in the
bottom, and the centre tube inside the catalyst, figure 4. At first the
process gas flows downwards through the catalyst before it enters the
center tube at the bottom of the reformer tube. The centre tube runs
through the whole catalyst. This make the TBR more energy efficient
since the hot exiting product gas heats the catalyst bed on its way out
and therefore decrease the need for external heat. [19]
2.3.1.2. Heat exchange reforming
The two, modern, heat exchange reformers available from Topsoe are
the Haldor Topsoe Convective Reformer (HTCR) and the Haldor Topsoe
Exchange Reformer (HTER). [15] The HTCR, figure 6, consists of a vertical
section containing the tube bundle and a horizontal combusting chamber.
The tube assembly follows the same principle as the bayonet reformer
explained above with a tube for exiting product gas inside the reformer
Figure 4 Principle of Topsoe Bayonet Reformer [19]
Figure 3 a) Inlet header; b) Flue-gas duct; c) Wall burners; d) Outlet manifold [20]
6
tube. Outside the reformer tube there is another tube for the flue gas, figure 5. Heat is then
transferred from the hot flue gas to the catalyst. [21] HTCR is ideal for smaller plants with a capacity
up to 30 000 Nm3/h of pure hydrogen and where no steam export is needed. [16]
The HTER is used either as a revamp in an
existing plant or as a way of making a new
plant more efficient e.g. lower the overall
natural gas consumption and the steam
production. The principle of the HTER is that
two reformers are used. The hot effluent gas
from the main reformer is used as heat
source in the second one. The tubes in the
HTER are double with the heating gas passing
on the inside and with catalyst on both sides
of the tube, figure 7. [15]
2.4. Water gas shift reaction In order to lower the concentration of carbon
monoxide in the gas and increase the yield of
hydrogen, a water gas shift (WGS) reactor is usually
used. The WGS reactor often consists of two units,
one that operates at a higher temperature (HTS)
and one that operates at a lower temperature (LTS). The HTS operates between 350 °C and 500 °C
and its catalyst could consist of an iron/chrome oxide. The LTS operates between 180 °C and 250 °C
Figure 7 Principle of HTER [22]
Figure 5 Principle of HTCR tube [15] Figure 6 Cross-cut of a HTCR [21]
7
and can use a copper/zinc oxide as catalyst. Even magnesium oxide and caesium oxide can be used as
catalysts in LTS.
Figure 8 Reaction mechanism for WGS reaction on CeO2. [23]
The reaction is:
𝐶𝑂 + 𝐻2𝑂 → 𝐶𝑂2 + 𝐻2 Δ𝐻2980 = −41 𝑘𝐽/𝑚𝑜𝑙𝑒 (2.4.-1)
Sometimes only the HTS is used and not the LTS. The WGS can in some cases be replaces by a partial
oxidation reactor (POX). [3, 23]
Haldor Topsoe A/S has developed medium temperature catalysts to replace both the high and the
low temperature shift. They operate between 190 °C and 330 °C. The two different catalysts, Cu-Zn-
Cr and Cu-Zn-Al, are used at the same time in layers. [24]
2.5. Off-gas purification A typical dry composition of the off-gas from the WGS is 70 – 80 % H2, 15 – 25 % CO2, 3 – 6 % CH4, 1 –
3 % CO, and some trace amounts of N2. This stream has to be purified; in the industry four different
ways are used. They are adsorption, absorption, membranes and cryogenic processes.
Today pressure swing adsorption (PSA) is the most common purification unit in new hydrogen plants,
because of the high purity that easily can be reached and of its relatively low cost. If an absorption
unit, which scrubs with a weak base like potassium carbonate or an amine like ethanolamine, and a
methanation reactor is used the product gas contains 95 – 97 vol.% H2, 2 – 4 vol.% CH4 and 0 – 2
vol.% N2. This could be used in an ammonia process where the purity is not essential. But for the
same money you could get a PSA that gives a purity of 99.99 vol. % with a hydrogen recovery up to
90 % and if an even purer gas is needed the PSA can give purity higher than 99.9999 vol. % hydrogen,
but with poorer recovery.
Membrane processes use polymeric membranes with one or more layers of different polymers. The
active layer is usually a polysulfone. The hydrogen purity reached in this process varies between 70
and 99 % and the recovery is between 70 and 95 %.
Before the off-gas enters any of these purification processes it has to be cooled in order for the water
to condensate as much as possible. [25, 3]
8
2.5.1. Pressure swing adsorption
The idea of pressure swing adsorption is that when the pressure in the column is high the adsorbents
adsorb most of the impurities in the gas stream so that almost pure hydrogen comes out of the unit
and when the pressure is low these impurities will desorb and the adsorbent will be regenerated. In a
PSA two different adsorbents are used. First a layer with activated carbon that adsorbs water and
carbon dioxide and afterwards a layer of zeolite 5A that adsorbs methane, nitrogen, carbon
monoxide and the remaining carbon dioxide. The activated carbon typically, for a normal
composition of steam methane reforming off-gas (SMROG), makes up 76% of the column height and
the rest is zeolites.
Usually ten or eleven columns are used but this can vary, from four in a small plant up to 16 in the
biggest plants. The process is cyclic, the pressure in the column to be regenerated is reduced in
several steps. Each reduction of the pressure in the column corresponds to an equal increase of the
pressure in another column. The waste gas that normally contains a lot of combustible gases is sent
to the reformer to be combusted and generate heat for the reformer.
Figure 9 Schematic flow diagram for a PSA process. [25]
The pressure is then increased in a corresponding way to the decrease and the vessel is ready to be
used again. The columns are usually regenerated when the adsorbents are saturated to the extent of
a half. This is done in order to keep the high purity of the product gas, but also to assure that no
carbon dioxide will reach the zeolite, this by making sure that the carbon dioxide is completely
adsorbed on the activated carbon. This is necessary because carbon dioxide is strongly adsorbed on
the zeolites making it impossible for them to be regenerated by a pressure swing.
Sometimes the desorption gas is taken to membranes where hydrogen is recovered and sent back to
the PSA before the gas is combusted. [25, 26, 3]
9
3. Kinetics and theory of formation of side products
3.1. Shift reaction with the surface intermediate HCOOH The shift reaction of CO, with a Cu/ZnO catalyst, proceeds with an intermediate on the catalyst
surface. The intermediate can be described as a complex formed from one molecule of each of the
reactants in the shift reaction.
𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐶𝑂2 + 𝐻2 Shift reaction [27] (3.1.-1)
𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐻𝐶𝑂𝑂𝐻 Decarbonylation [28] (3.1.-2)
𝐶𝑂2 + 𝐻2 ⇄ 𝐻𝐶𝑂𝑂𝐻 Decarboxylation [28] (3.1.-3)
The carbonylation and decarboxylation of formic acid can be compared with the forward and reverse
shift reactions [28]. Both forward and reverse shift reactions are preceded by the catalyst surface
intermediate. The rates of the reactions are controlled by decomposition of the intermediate or by
desorption rates. [27] The rate equations for the shift equations correspond to the decomposition of
the intermediate complex as the rate determining. In this case it will be described with
decomposition of formic acid. [29]
The hydrogenation (backward carboxylation) can be described with a Langmuir-type kinetic
equation, equation 1 below describes a typical rate equation of the hydrogenation, the equation is
only valid at low pressures.
𝑟 =𝑘∞ exp −
𝐸
𝑅𝑇 𝑝𝐻𝐶𝑂𝑂𝐻
1+𝐾0 exp −∆𝐻
𝑅𝑇 𝑝𝐻𝐶𝑂𝑂𝐻
𝑚 𝑚𝑜𝑙𝑒/(𝑔 𝑠) [29] (3.1.-eq1)
Summarization of the shift reaction with the intermediate, both forward and reverse CO shift
reactions are controlled by a surface complex with the same stoichiometry as formic acid, as the
reaction below.
The kinetic of these reactions are readily available in the literature but it is important to understand
the theory around the complex intermediate, formic acid.
𝐶 + 𝐻2𝑂 ⇄ (𝐻2𝐶𝑂2) ⇄ 𝐶𝑂2 + 𝐻2 [27] (3.1.-4)
10
Figure 10 Potential energy diagram of the water-gas shift reaction on Cu(110). [30]
3.2. Ammonia The equations below are the classic rate equations for the ammonia synthesis.
𝑟 = 𝑘𝐴𝑝𝑁2
𝑃𝐻2 3
𝑃𝑁𝐻3 2
𝛼
− 𝑘𝐵 𝑝𝑁𝐻3
2
𝑝𝐻2 3
1−𝛼
(3.2.-eq1)
𝑟 =𝑘𝐴
′ 𝑝𝑁2−𝑘𝐵′ 𝑝𝑁𝐻3
2
𝑃𝐻23
1+𝐾𝑐 𝑃𝑁𝐻3
𝑃𝐻2
32
2𝛼 (3.2.-eq2)
Equation 3.2.-eq1, Temkin-Pyzhev´s, equation is useful assuming that the catalyst surface coverage,
by atomic nitrogen is high and equation 3.2.-eq3, Ozaki´s, when atomic nitrogen coverage is low.
Equation 3.2.-eq2 assumes that nitrogen dissociative adsorption is rate determining. [31, 32]
Equilibrium equations for ammonia are readily available in the desired interval for temperature and
pressure. If the equilibrium constant for the interval of interest cannot be found an alternative is to
interpolate from the available data. In the studied literature there are no indications that ammonia is
formed on the current catalyst. Its formation is only catalyzed by Ce, Ti, Zr, V, Cr, Mo, W, U, Mn, Re,
Fe, Ni, Co and Ru. [33].
The ammonia out from the reformer is at equilibrium, volume ppm wet basis, catalyzed with Ni. [34]
The equilibrium equation for ammonia:
𝑙𝑜𝑔10 𝐾𝑝
𝐾𝑝∗ =
0.1191849
𝑇+
25122730
𝑇4 +38.76816
𝑇2 𝑥𝑖𝐴0𝑖
1/2+
64.49429
𝑇2 𝑥𝑖𝐴0𝑖
1/2
2 𝑝
(3.2.-eq3)
𝑙𝑜𝑔10 𝐾𝑝∗ = −2.691122 ∙ 𝑙𝑜𝑔10 𝑇 − 5.519265 ∙ 10−5𝑇 + 1.848863 ∙ 10−7𝑇2 +
2001 .6
𝑇+ 2.6899 (3.2.-eq4)
11
Pressure atm, 𝜈𝑖 reaction coefficient for substance i and constants in table 5 below:
Table 5 Constants in eq. 3 above. [35]
Gas A0 B0 c
Hydrogen 0.1975 0.02096 0.0504∙ 104
Nitrogen 1.3445 0.05046 4.20∙ 104
Ammonia 2.3930 0.03415 476.87∙ 104
3.3. Methanol Methanol can be produced in the shift reactor in two different ways, either from carbon dioxide or
from carbon monoxide:
𝐶𝑂 + 2𝐻2 ⇄ 𝐶𝐻3𝑂𝐻 (3.3.-1)
𝐶𝑂2 + 3𝐻2 ⇄ 𝐶𝐻3𝑂𝐻 + 𝐻2𝑂 (3.3.-2)
The copper/zinc catalyst in the shift reactor is practically the same one that is used in a methanol
reactor when methanol is produced from synthesis gas. This means that the shift converter is a good
methanol converter as well. The big difference is the pressure; methanol synthesis is favoured by
high pressure, why the pressure in the shift reactor is of great importance for the unwanted
methanol side formation. [34]
Three different reaction mechanisms of the reaction with carbon monoxide are shown below. Which
of these that is most important is not known. [36]
However, it can be argued that the formation of methanol from carbon monoxide is of minor
importance and can be neglected in calculations. [37]
Because of the fact that methanol synthesis is usually
operated at higher pressures, compared to those in
the shift reactor, most literature studies describe the
kinetics at these conditions.
Below are some kinetic expressions that are of
interest.
Leonov et al. have done experiments with a copper-
zinc oxide-alumina catalyst at low pressures and
temperatures between 220 and 260 °C. [38] The
kinetic expression is:
𝑟 = 𝑘 𝑝𝐶𝐻
0.5 ∙𝑝𝐻2
𝑝𝐶𝐻3𝑂𝐻0.66 −
𝑝𝐶𝐻3𝑂𝐻0.34
𝑝𝐶𝑂0.5∙𝑝𝐻2
∙𝐾𝑒𝑞 (3.3.-eq1)
However the k-value is not reported.
Klier et al. have created a more complex model for
Cu/ZnO catalysts at low temperatures. [38]
𝑟 = 𝑐𝑜𝑛𝑠𝑡 ∙ 𝐴𝑜𝑥𝑚 ∙
𝑝𝐶𝑂 ∙𝑝𝐻22 −𝑝𝑀𝑒𝑂𝐻 /𝐾𝑒𝑞
𝐹+𝐾𝐶𝑂2 ∙𝑝𝐶𝑂2 𝑛 + 𝑘′ ∙ 𝑝𝐶𝑂2
− 𝐾𝑒𝑞′
−1∙ 𝑝𝑀𝑒𝑂𝐻 ∙ 𝑝𝐻2𝑂/𝑝𝐻2
3 (3.3.-eq2)
Figure 11 Different reaction mechanisms of the formation of methanol. [36]
12
𝐴𝑜𝑥 = 𝐴0 ∙𝐾′∙ 𝑝𝐶𝑂2 /𝑝𝐶𝑂
1+𝐾′∙ 𝑝𝐶𝑂2/𝑝𝐶𝑂
(3.3.-eq3)
Table 6 Expressions in Klier et al´s equation.
𝐾𝑝 = 3.27 ∙ 10−13 ∙ exp(11678/𝑇) 𝐾𝑝′ = 3.826 ∙ 10−11 ∙ exp(6851/𝑇)
𝐾𝛾 = 1 − 𝐴1 ∙ 𝑃 𝐾𝛾′ = 1 − 𝐴1 ∙ 𝑃 ∙ 1 − 𝐴2 ∙ 𝑃
𝐴1 = 1.95 ∙ 10−4 ∙ 𝑒𝑥𝑝 1703/𝑇 𝐴2 = 4.24 ∙ 10−4 ∙ 𝑒𝑥𝑝 1107/𝑇
𝐾𝑒𝑞 = 𝐾𝑝/𝐾𝛾 𝐾𝑒𝑞′ = 𝐾𝑝
′ /𝐾𝛾′
𝐹 = 1 + 𝐾𝐻2∙ 𝑝𝐻2
+ 𝐾𝐶𝑂 ∙ 𝑝𝐶𝑂 𝑚 = 3 𝑛 = 3
Table 7 Constants in Klier et al´s equation.
Case Temp ℃ 𝑲𝑪𝑶 𝑲𝑯𝟐 𝑲𝑪𝑶𝟐 𝑲′ 𝒌𝑨𝟎
𝟑 𝒌′ 𝑲𝒆𝒒 𝑲′𝒆𝒒
I 225 12.52 1.77 39.62 158.2 1.064 2.18 9.034 9.237 235 8.58 1.40 21.52 125.4 1.253 2.70 5.409 6.573 250 5.00 1.00 9.00 90.0 1.584 3.75 2.625 4.095
II 225 12.52 1.77 98.51 18.1 0.898 2.18 9.034 9.237 235 8.58 1.40 60.0 16.1 1.135 2.70 5.409 6.573 250 5.00 1.00 30.0 13.8 1.584 3.75 2.625 4.095
III 225 12.52 1.77 19.8 79.1 0.088 2.18 9.034 9.237 235 8.58 1.40 10.8 62.7 0.120 2.70 5.409 6.573 250 5.00 1.00 4.5 45.0 0.195 3.75 2.625 4.095
Also in this model the constant “const” is unreported.
Jennings et al. have created a model for the equilibrium constant built on the activity of the
components. N is for the number of moles. [39]
𝐾𝑠𝑦𝑛𝑡 𝑒𝑠𝑖𝑠 = 𝑁𝐶𝐻3𝑂𝐻 ∙𝑁𝑇
2
𝑁𝐶𝑂 ∙𝑁𝐻22 ∙𝑃2 ∙
𝛾𝐶𝐻3𝑂𝐻
𝛾𝐶𝑂 ∙𝛾𝐻22 (3.3.-eq4)
𝐾𝑠𝑖𝑓𝑡 = 𝑁𝐶𝑂 ∙𝑁𝐻2𝑂
𝑁𝐶𝑂2 ∙𝑁𝐻2
∙ 𝛾𝐶𝑂 ∙𝛾𝐻2𝑂
𝛾𝐶𝑂2 ∙𝛾𝐻2
(3.3.-eq5)
𝑁𝑇 = 𝑁 (3.3.-eq6)
Skrzypek et al. have made a model on the assumption that only carbon dioxide and not carbon
monoxide is a part of the reaction. The model shown below is for 3-9 MPa, 460-550 K and a
Cu/ZnO/Al2O3 catalyst. [37]
𝐶𝑂2 + 3𝐻2 ⇄ 𝐶𝐻3𝑂𝐻 + 𝐻2𝑂 (r1, 3.3.-eq7) (3.1.-3)
𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐶𝑂2 + 𝐻2 (r2, 3.3.-eq8) (3.1.-4)
𝑟1 = 𝑘1 ∙ 𝐾𝐻2
2 ∙ 𝐾𝐶𝑂2∙
𝑝𝐻22 ∙𝑝𝐶𝑂2 − 1/𝐾𝑝1 𝑝𝐶𝐻3𝑂𝐻 ∙𝑝𝐻2𝑂/𝑝𝐻2
1+𝐾𝐻2 ∙𝑝𝐻2 +𝐾𝐶𝑂2 ∙𝑝𝐶𝑂2 +𝐾𝐶𝐻3𝑂𝐻 ∙𝑝𝐶𝐻3𝑂𝐻 +𝐾𝐻2𝑂 ∙𝑝𝐻2𝑂+𝐾𝐶𝑂 ∙𝑝𝐶𝑂 3 (3.3.-eq7)
𝑟2 = 𝑘2 ∙ 𝐾𝐻2∙ 𝐾𝐶𝑂2
∙ 𝑝𝐻2 ∙𝑝𝐶𝑂2 − 1/𝐾𝑝2 𝑝𝐶𝑂 ∙𝑝𝐻2𝑂
1+𝐾𝐻2∙𝑝𝐻2
+𝐾𝐶𝑂2∙𝑝𝐶𝑂2
+𝐾𝐶𝐻3𝑂𝐻 ∙𝑝𝐶𝐻3𝑂𝐻 +𝐾𝐻2𝑂 ∙𝑝𝐻2𝑂+𝐾𝐶𝑂 ∙𝑝𝐶𝑂 2 (3.3.-eq8)
𝑘𝑖 = 𝑘𝑖0 ∙ 𝑒𝑥𝑝 −𝐸𝑖
𝑅∙𝑇 (3.3.-eq9)
13
𝐾𝑖 = 𝐾𝑖0 ∙ 𝑒𝑥𝑝 −∆𝐻𝑖
𝑅∙𝑇 (3.3.-eq10)
Table 8 Constants in the equation above.
𝑘10 = 3 ∙ 109 𝑘𝑚𝑜𝑙/(𝑘𝑔 𝑐𝑎𝑡 ∙ ) 𝑘20 = 2.5 ∙ 109 𝑘𝑚𝑜𝑙/(𝑘𝑔 𝑐𝑎𝑡 ∙ )
𝐸1 = 104.7 𝑘𝐽/𝑚𝑜𝑙 𝐸2 = 104.7 𝑘𝐽/𝑚𝑜𝑙
Table 9 Constants in the equation above.
𝐾𝐻2 𝐾𝐶𝑂2
𝐾𝐶𝐻3𝑂𝐻
𝐾10 = 0.14 ∙ 10−8 𝑎𝑡𝑚−1 𝐾20 = 0.44 ∙ 10−8 𝑎𝑡𝑚−1 𝐾30 = 0.11 ∙ 10−9 𝑎𝑡𝑚−1
∆𝐻1 = −75.4 𝑘𝐽/𝑚𝑜𝑙 ∆𝐻2 = −75.4 𝑘𝐽/𝑚𝑜𝑙 ∆𝐻3 = −29.3 𝑘𝐽/𝑚𝑜𝑙
𝐾𝐻2𝑂 𝐾𝐶𝑂
𝐾40 = 0.35 ∙ 10−8 𝑎𝑡𝑚−1 𝐾50 = 0.50 ∙ 10−10 𝑎𝑡𝑚−1
∆𝐻4 = −75.4 𝑘𝐽/𝑚𝑜𝑙 ∆𝐻5 = −75.4 𝑘𝐽/𝑚𝑜𝑙
Ostrovskii et al. have made a model for Cu-Zn catalysts, 0.1-25 MPa and 470-570 K. However, the two
constants k3 and k4 are undefined. [40]
𝑟 = 𝐴/𝐵 (3.3.-eq11)
𝐴 = 𝑘4 ∙ 𝐾2 + 𝑘3 ∙ 𝐾1 ∙ 𝑃𝐻2∙ 𝑃𝐶𝑂2
∙ 1 − 𝑥 (3.3.-eq12)
𝐵 = 1 + 𝐾1 ∙ 𝑃𝐶𝑂2+ 𝐾2 ∙ 𝑃𝐻2
+ 𝑃𝐶𝐻3𝑂𝐻/𝐾6 ∙ 𝑃𝐻2 + 𝑃𝐶𝐻3𝑂𝐻 ∙ 𝑃𝐻2𝑂 ∙
𝑃𝐶𝑂2
𝐾52 ∙ 𝐾6
2 ∙ 𝐾7 ∙
𝑃𝐻2 0.5
+ 𝑃𝐻2∙ 𝑃𝐻2𝑂 ∙ 𝑃𝐶𝑂2
/𝐾7 ∙ 𝑃𝐶𝐻3𝑂𝐻 0.5
(3.3.-eq13)
𝑥 = 𝑃𝐶𝐻3𝑂𝐻0.5 ∙ 𝑃𝐻2𝑂
0.5 /𝐾0.5 ∙ 𝑃𝐻2
1.5 ∙ 𝑃𝐶𝑂2
0.5 (3.3.-eq14)
Table 10 Constants in Ostrovskii´s equation.
Constant Value Dimension
𝑲𝟏 4.91 ∙ 10−4 ∙ 103085/𝑇 𝑀𝑃𝑎−1 𝑲𝟐 2.1 ∙ 10−4 ∙ 102348/𝑇 𝑀𝑃𝑎−1 𝑲𝟔 0.00288 ∙ 10−501/𝑇 𝑀𝑃𝑎−1
𝑲𝟕𝟎.𝟓 4.3 ∙ 106 ∙ 10−4940/𝑇 𝑀𝑃𝑎
𝑲𝟓 ∙ 𝑲𝟔 ∙ 𝑲𝟕𝟎.𝟓 2.33 ∙ 104 ∙ 104290/𝑇 𝑀𝑃𝑎
3.4. Methylamines Methylamines are produced on a shift catalyst when ammonia and methanol is present. [41] Three
different methylamines can occur, monomethylamine (MMA), dimethylamine (DMA) and
trimethylamine (TMA), which all are unwanted toxic by-products. The reactions are shown below:
2𝐶𝐻3𝑁𝐻2 ⇄ 𝐶𝐻3 2𝑁𝐻 + 𝑁𝐻3 [42] (3.4.-1)
2 𝐶𝐻3 2𝑁𝐻 ⇄ 𝐶𝐻3 3𝑁 + 𝐶𝐻3𝑁𝐻2 [42] (3.4.-2)
𝐶𝐻3 2𝑁𝐻 + 𝐶𝐻3𝑁𝐻2 ⇄ 𝐶𝐻3 3𝑁 + 𝑁𝐻3 [42] (3.4.-3)
𝑁𝐻3 + 𝐶𝐻3𝑂𝐻 ⇄ 𝐶𝐻3𝑁𝐻2 + 𝐻2𝑂 [43] (3.4.-4)
𝐶𝐻3 2𝑁𝐻 + 𝐶𝐻3𝑂𝐻 ⇄ 𝐶𝐻3 3𝑁 + 𝐻2𝑂 [43] (3.4.-5)
14
The formation of TMA is thermodynamically favoured, but the amount of each amine is strongly
dependent on the catalyst. Different catalysts have different selectivity for each amine. [44]
The equilibrium constants for the distribution between the three methylamines are presented below:
[42]
877.0)/4.1028()log( 1 KTK (3.4.-eq1)
792.0)/8.687()log( 2 KTK (3.4.-eq2)
673.1)/1.1725()log( 3 KTK (3.4.-eq3)
One kinetic expression produced by Mitchell et al. is as follow: [42]
34
2
11 NHDMAMMA CCkCkr (3.4.-eq4)
TMAMMADMA CCkCkr 5
2
22 (3.4.-eq5)
3633 NHTMADMAMMA CCkCCkr (3.4.-eq6)
114 / Kkk (3.4.-eq7)
225 / Kkk (3.4.-eq8)
336 / Kkk (3.4.-eq9)
)/29451exp(1088.1 9
1 KTRk )/(9872.1 KmolgcalR
)/24448exp(1077.9 7
2 KTRk )]/([ 2
catgmolgLk
)/34063exp(1033.3 11
3 KTRk )]/([ catgsmolgr
Laboratory experiments have shown that the formation of methylamine in the shift reactor is
proportional to the formation of methanol. This means that the higher the formation rate of
methanol, the higher the methylamine rate of formation. [34]
4. Influence of contaminants
4.1. Ammonia When water contains ammonia it readily attacks copper and copper bearing alloys. It attacks the
metals by forming the copper-ammonium ion whereupon they dissolve. The resulting corrosion leads
to deposits on boiler heat transfer surfaces and reduces efficiency and reliability. [45] This means
that more expensive material has to be used.
4.2. Carbon dioxide The principal effect of dissolved carbon dioxide is corrosion. When the gas dissolves in water it
produces carbonic acid which lowers the pH and makes it corrosive.
𝐶𝑂2 + 𝐻2𝑂 ⇄ 𝐻2𝐶𝑂3 ⇄ 𝐻+ + 𝐻𝐶𝑂3− (4.2.-1)
The lowering of pH also enhances the corrosive effects of oxygen on iron. [46]
𝐹𝑒 + ½𝑂2 + 2𝐻+ ⇄ 𝐹𝑒2+ + 𝐻2𝑂 (4.2.-1)
15
4.3. Carboxylic acids Also the carboxylic acids cause corrosion due to lowering of the pH. If there is methanol present in
the water when the carbon monoxide dissolves, acetic acid can be formed. [47]
𝐶𝑂 + 𝐶𝐻3𝑂𝐻 ⇄ 𝐶𝐻3𝐶𝑂𝑂𝐻 ⇄ 𝐶𝐻3𝐶𝑂𝑂− + 𝐻+ (4.3.-1)
4.4. Oxygen The electrochemical process that describes oxygen attack on iron, called oxygen pitting, is as follows:
𝐴𝑛𝑜𝑑𝑒: 𝐹𝑒 → 𝐹𝑒2+ + 2𝑒− (4.4.-1)
𝐶𝑎𝑡𝑜𝑑𝑒: ½𝑂2 + 𝐻2𝑂 + 2𝑒− → 2𝑂𝐻− (4.4.-2)
𝑂𝑣𝑒𝑟𝑎𝑙𝑙: 𝐹𝑒 + ½𝑂2 + 𝐻2𝑂 → 𝐹𝑒 𝑂𝐻 2 (4.4.-3)
Oxygen is highly corrosive when present in hot water because the high temperature provides enough
energy to accelerate reactions at the metal surfaces, resulting in rapid and severe corrosion. This
causes serious problems even at low concentrations. Oxygen corrosion can result in rapid failure
because pits can penetrate deep into the metal and is commonly a problem in boiler tubes,
economizers and condensate lines. It can also lead to problems with iron deposits from iron oxide
that the corrosion produces. [48]
5. Condensate treatment
5.1. Biological nitrogen treatment One way to reduce nitrogen in water is to treat the water biologically. This will be done by bacteria
which oxidize ammonium (𝑁𝐻4+) to nitrate (𝑁𝑂3
−) in an aerobic environment. This process is called
nitrification. The oxygen consumption in the process is about five times the ammonia content and it
is very temperature dependent. At lower temperatures (about 5oC) the digestion rate will decrease.
To compensate this, the bacteria concentration could be increased. [49]
After the aerobic process the nitrate will be converted to 𝐶𝑂2, 𝐻2𝑂 and 𝑁2. This is called
denitrification. In this step the bacteria “breath” the oxygen bound in the nitrate. This is called anoxic
digestion. The nitrogen that is released will leave the water as nitrogen gas. The anoxic digestion is
not as fast as the aerobic process and it is much more dependent on the temperature. The bacteria
also need a carbon source, for example methanol. It is also very important that there is no oxygen
present in the process since the bacteria rather use free oxygen than nitrogen bound oxygen for
breathing. [49]
5.1.1. Nitrification
Nitrification is when ammonia is converted to nitrate. First the ammonia must be converted to
ammonium which will be done by decreasing the pH. The ammonium is then converted to nitrate in
a two step reaction by autotrophic bacteria according to the following reaction:
𝑁𝐻4+ + 1.5𝑂2 → 𝑁𝑂2
− + 2𝐻+ + 𝐻2 (5.1.1.-1)
𝑁𝑂2− + 0.5𝑂2 → 𝑁𝑂3
− (5.1.1.-2)
16
As can be seen acid is formed, this will decrease the pH. A lower pH will inhibit the bacteria activity
which will lead to a decreased reaction rate.
As mentioned earlier the nitrification rate is dependent on the temperature. If the temperature
reaches 4°C the bacteria work very slowly and the residence time will increase to 20 days. This will
lead to the need of very big cisterns resulting in high costs. [49]
5.1.2. Denitrification
In the denitrification step the microorganisms live in an oxygen free environment. The
microorganism oxidizes organic material which can be methanol and reduces the nitrite. This can be
described by the following chemical equation:
2𝑁𝑂3− + 𝐻+ + 𝑜𝑟𝑔𝑎𝑛𝑖𝑐 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙 → 𝑁2 + 𝐻𝐶𝑂3
− (5.1.2.-1)
The nitrogen gas produced is vented to the atmosphere. The optimal pH for this process is around 7-
9.The amount of methanol that is needed in the denitrification step is 2.8 g methanol/g reduced
nitrogen. [49]
5.1.3. Other treatment processes
An alternative to the traditional denitrification processes are the Anammox process. This process
convert ammonium to nitrogen gas and uses nitrite as an electron acceptor. The process is
autotrophic which means that there is no need for adding methanol. The Anammox process can be
combined with a nitrification step. This means that a lower level of ammonium conversion is needed.
The remaining ammonia and the nitrite will be converted into nitrogen gas in the Anammox process.
This leads to a decrease in oxygen demand in the nitrification step and a reduction of the process
costs. The degree of purification for this process, when a fluidized-bed reactor was used, is 82 % for
the ammonia and >99 % for the nitrite. [50]
One way to improve an already existing water treatment plant is to use the SHARON process. This
process contains a batch reactor which runs first in a nitrification phase and then switches to a
denitrification phase. The process temperature must be around 35 °C and then the nitrogen
concentration in the effluent will vary from <5 to about 100 mg/l depending on the residence time.
[51, 52]
5.2. Removal of methanol with a membrane bioreactor In a membrane bioreactor microorganisms consume some substances, in this case methanol. The
biomass is then retained by the membrane [52]. A bioreactor is capable of very high removal
efficiencies, above 99 %. One report states that the methanol concentration was reduced from 100
to less than 0.5 mg/L. Optimal operating conditions were; 60 °C and pH around 6. [53] The same
results were reported from other trials [54]. The microorganisms need six weeks to acclimatize and
reach steady state. The process was operated as a fed-batch process, in which the waste condensate
and nutrition were fed periodically. [53]
5.3. Stripping Stripping is a well known and frequently used method for removal of volatile organic compounds and
dissolved gases from wastewater. [55, 56]
17
There are two major types of stripping processes, air and steam stripping. The basic idea is the same,
the liquid wastewater stream is fed to the top of the stripping column and flows counter currently to
the gas stream. When the liquid comes in contact with the gas the dissolved volatile gaseous
impurities in the liquid are transferred to the gas. The gas is either air or steam and therefore called
air and steam stripping. [52] The stripper column is basically a distillation column. It is either a
packed tower or a tray column, e.g. bubble-cap. [56, 57, 58]
Air and steam strippers display different removal efficiencies for the compound of interest in this
report. Both experimental research and real life results from industry shows that ammonia removal
by air stripping is effective. Trials in a 20 tray bubble-cap column give stripping efficiencies of 98 – 99
% at optimal operating conditions, air to feed ratio about 2.5, pH 10.8, feed temperature 30 – 35 °C.
The feed concentration varied between 640 to 720 mg/kg and the bottom concentration between 2
and 13 mg/kg. The feed also contained methanol at concentrations varying between 700 and 800
mg/kg. Concentration varied extensively in the bottom, from 40 to 180 mg/kg at optimal operating
conditions, as above, thus giving a removal efficiency of 74 to 95 %. [56] Other studies shows
ammonia removal efficiencies of above 90 %, feed containing about 4500 mg/L ammonia. [59]
The shortcomings of air stripping of methanol are confirmed in a report on industrial condensate
strippers. Total removal efficiencies ranging from 0 to 25 % were reported. Steam stripping on the
other hand shows greater potential for methanol removal. This was investigated in the same report.
Removal efficiencies of up to 99 % were reported. The feed concentration ranged between 130 and
8000 ppmw with a median of 4800 ppmw. [60] Another report on industrial steam stripping gives a
removal of methanol from 5226 to 100 mg/kg. Notable is that the performance of the actual process
was considered not satisfactory. The aim of the report was to create a simulator of the process in
order to find improvements. The results indicate that the methanol concentration could be reduced
to 50 mg/kg. [61] A third report indicates removal efficiencies above 98 % for steam stripping of
methanol in the industry. [62]
Steam stripping displays the same efficiency as air stripping when it comes to ammonia removal. [57,
61, 63] It also has the capability to drastically remove carbon dioxide in the stream, from 1500 to 1
ppmw. Notable is that these results were achieved at pH level far too low for effective ammonia
removal, as described below. [63]
The most important design factor in steam strippers in general is the steam to feed ratio. In this
particular stripper configuration pH is of equal importance.
There are a number of equations of interest in the system CO2/NH3 /H2O:
𝐻2𝑂 ⇄ 𝐻+ + 𝑂𝐻− (5.3.-1)
𝑁𝐻3 + 𝐻2𝑂 ⇄ 𝑁𝐻4+ + 𝑂𝐻− (5.3.-2)
𝐶𝑂2 + 𝑂𝐻− ⇄ 𝐻𝐶𝑂3− (5.3.-3)
𝐶𝑂2 + 𝐻2𝑂 ⇄ 𝐻𝐶𝑂3− + 𝐻+ (5.3.-4)
𝐻𝐶𝑂3− ⇄ 𝐶𝑂3
2− + 𝐻+ (5.3.-5)
𝑁𝐻3 + 𝐶𝑂2 ⇄ 𝑁𝐻2𝐶𝑂𝑂− + 𝐻+ (5.3.-6)
18
Reaction (5.3.-1) is instantaneous. The same goes for reaction (5.3.-2) the equilibrium between
ammonia and its ionic form, ammonium ion. Carbon dioxide in water forms the carbonate ion via
reactions (5.3.-3) – (5.3.-5), (5.3.-3) is fast, (5.3.-4) slow and (5.3.-5) instantaneous. The formation of
the ammonium carbamate ion (5.3.-6) is fast. [64]
Ammonia and carbon dioxide can only be stripped from the water solution when in their molecular
form [65]. Which form that is present is controlled by the equilibrium equations (5.3.-1) – (5.3.-6).
They are both dependent of temperature; high temperature reduces the solubility of the gases in
water and thus increases the removal. [64] The most important factor affecting if molecular or ion
form is present, and thus the removal efficiency, is pH. A higher pH, that is a higher concentration of
the hydroxyl ion, shifts reaction (5.3.-2) towards NH3 and thus favours molecular ammonia over the
ammonium ion. The opposite can be said about carbon dioxide, a low pH favours carbon dioxide over
carbonate. In a system of both ammonia and carbon dioxide the removal of one component affects
the other. If we look at the removal of carbon dioxide, equation (5.3.-3), we neglect equation (5.3.-4)
since equation (5.3.-3) is much faster, it can be seen that this leads to an increase of the hydroxyl ion
since the reaction is shifted to the left. An increase of the hydroxyl ion shifts reaction (5.3.-2) to the
right and thus more ammonia can be removed. Removal of either ammonia or carbon dioxide
directly affects the formation of the other via reaction (5.3.-6).
A method of calculating the required pH is:
𝑝𝐻 = − log 𝐾𝑤
𝐾𝑏
𝐶𝑜𝑢𝑡
𝐶𝑖𝑛 −𝐶𝑜𝑢𝑡 (5.3.-eq1)
Were 𝐾𝑤 is the ion product for water, 𝐾𝑏 the equilibrium constant for the NH3 – H2O system, 𝐶𝑜𝑢𝑡 is
the ammonia concentration in the effluent and 𝐶𝑖𝑛 is the ammonia concentration in the influent. [65]
Investigations show that an appropriate pH should at least be above 10, preferably close to 11 [63,
64, 65]. Another report states that pH should be above 11 [59]. Ways to increase pH includes adding
NaOH [63] and lime [60]. As a comparison data at a pH of 7.2 is presented in table 11. As seen most
ammonia is present as ammonium, interesting is also that carbonate dominates over carbon dioxide.
Table 11 Equilibrium composition at 100 °C and pH 7.2 [63]
NH3 present as % CO2 present as %
NH3 34.7 CO2 7.8
NH4+ 64.4 HCO3
- 87.9
NH2COO- 0.9 CO32- 0.1
NH2COO- 4.0
A report shows that carbon dioxide is the dominant of the two, CO2 and NH3. In the presence of high
concentrations of carbon dioxide the ammonia removal is dependent on the removal rate of carbon
dioxide. To obtain a large driving force for removal of carbon dioxide the concentration of carbon
dioxide should be kept low in the gas used for stripping. [64] This can be achieved if steam is used.
As mentioned before steam to liquid ratio is of great importance in stripping. Simulations of the
system, methanol – water show that the steam to feed ratio should be above 0.2 (kg/kg) to ensure
optimal methanol removal [62]. When it comes to ammonia, it is shown that increasing the steam to
feed ratio greatly affects the removal capacity. Stem to feed ratios as high as 0.3 (kg/kg) is preferred.
A thing that is important in this discussion is the number of trays in the column. By increasing the
19
number of trays almost the same effect as increasing the steam to feed ratio is experienced. The
effect wears off drastically at approximately ten trays; it is therefore not necessary to increase the
number of trays above 15 to 20. The same argument regarding steam to feed ratio is also valid for
carbon dioxide. [63] One conclusion to be drawn is that in order to get the best removal possibly the
steam to feed ratio should be kept high, about 0.3 (kg/kg), and the number of trays should be close
to 20.
To control the water content in the overhead vapour the top temperature is the key variable [57]. A
high temperature increases the amount of liquids present in the vapour. This can increase the cost of
stripping since the heat recovered in the overhead condenser, if such is present, decreases. The
overhead temperature also affects the removal efficiency. In the case of methanol removal the
temperature should be kept above 80 °C. [60] Another important thing concerning temperature is
the feed inlet temperature. This should be as close to the boiling point as possible otherwise a
portion of the steam is need for heating the feed. [60] A common way to do this is by heat
exchanging the feed with the effluent condensate stream [57, 60, 63].
There are two different designs available for the steam stripper which greatly affects investment and
productions costs. The steam can either be supplied to the bottom of the column as fresh steam or
via a reboiler, which vaporises the liquid, as in a conventional distillation column [63]. One drawback
of the former is the obvious need of fresh steam.
Something to keep in mind is that the presence of organic acids affects the removal of ammonium.
The acids bind the ammonia and thus have to be neutralized. [63]
5.4. Deaeration Dissolved oxygen and carbon dioxide in boiler feed water systems will cause corrosion and the
corrosion increases with the temperature. It is therefore very important to reduce the oxygen and
carbon dioxide content to as near zero as possible. [66]
To separate dissolved gases like oxygen, carbon dioxide and ammonia from water deaeration could
be used. There are two different ways to separate the dissolved gases. One way is to reduce the
pressure by vacuum and the dissolved gases will be vented. This technique is often used in food
industry where high temperature is undesirable. This technique is however not suited for carbon
dioxide removal. The other way to achieve separation is by introducing a new gas to the system. The
new gas that is introduced is often steam. This technique is called pressure deaeration and is very
common in boiler feed water preparation. [67]
5.4.1. Pressure Deaerators
In a pressure deaerator the incoming water is sprayed onto trays in the vessel. The vessel contains an
atmosphere of steam and water. [67] The contaminated water meets the hot steam which increases
the water temperature to a few degrees above the saturation temperature of the steam, around 105
°C. [66] This leads to a reduced partial pressure of the dissolved gases and a very good deaeration is
achieved. The oxygen and the carbon dioxide in the water are released to the steam. 97 – 98 % of the
oxygen can be released in this way. The remaining oxygen must be scrubbed thoroughly with steam
to reach acceptable limits. Often the removal is not good enough and oxygen scavenger must be
added. The deaerated water is collected in the feed water storage where a steam blanket protects
20
the water from contamination. The deaerated water can then leave the vessel and be stored in a
bigger vessel before it is transferred to the boiler. [67]
The carbon dioxide content in the atmosphere is normally about 0.04 % and this gives the surface
water a solved carbon dioxide concentration of 10 mg/l. [67] About 75 – 95 % of the carbon dioxide is
released in the deaeration. The remaining 5 –25 % cannot be released because the steam and
condensate makeup lowers the pH which shifts the carbonate/bicarbonate equilibrium. This will
result in more bicarbonate ions according to the equation below. [68]
𝐻2𝑂 + 𝐶𝑂2 ⇄ 𝐻2𝐶𝑂3 ⇄ 𝐻+ + 𝐻𝐶𝑂3− ⇄ 𝐻+ + 𝐶𝑂3
2− (5.4.1.-1)
At higher temperatures the bicarbonate and 80 % of the carbonates will be transformed to carbon
dioxide. This leads to an increasing of alkaline hydroxide which raises the pH from 7 into 10. This will
result in a production of carbon dioxide in the steam and carbonic acid will be formed at
condensation which decreases the pH. The feed water will neutralize the makeup water in the
deaerator and reducing pH to about 8. The carbon dioxide will be removed at the trays in the
deaerator. [68]
To remove the last influence of the carbon dioxide, neutralizing amines can be injected. These
amines can be cyclohexylamine, benzylamine and morpholine. The amines will also vaporize in the
boiler and leave with the steam. When the steam condensates, carbonic acid is produced from the
carbon dioxide and the amines will neutralize it. The injection of these amines will result in an
increase of the pH and lessen the damage on the boiler and its equipment. The injected amount must
be in proportion to the amount of free carbon dioxide. [68]
If the carbon dioxide content is under 5 ppm it can be removed either with ion exchange treatment
or chemical treatment by using caustic soda. [66]
5.5. Ion exchange The principle of ion exchange is that ions in a solution are adsorbed on a solid material and are
replaced by other ions equivalently charged released by the ion exchanger material. Typical ion
exchanger materials are polymer gel resins and zeolites. [69]
Ion exchange with zeolites, as ion exchanger material, has been proven able to remove ammonia at
low concentrations (2-10 ppm). [70] When ammonia dissolves in water a significant amount reacts
with the water forming ammonium ions.
𝑁𝐻3 + 𝐻2𝑂 ⇄ 𝑁𝐻4+ + 𝑂𝐻− (5.5.-1)
In an ion exchange unit the ammonium ion could be replaced by another cat ion. Typical exchange
ions are 𝐻+, 𝑁𝑎+, 𝐾+, 𝑀𝑔2+ and 𝐶𝑎2+. The ammonia uptake is very dependent on both the zeolite
used and what replacement ion that specific zeolite is prepared with. Possible reactions in the ion
exchanger are as follows: [70]
𝑁𝐻4+(𝑎𝑞) + 𝑐𝑎𝑡𝑖𝑜𝑛+ − 𝑧𝑒𝑜𝑙𝑖𝑡𝑒 ⇄ 𝑐𝑎𝑡𝑖𝑜𝑛+(𝑎𝑞) + 𝑁𝐻4
+ − 𝑧𝑒𝑜𝑙𝑖𝑡𝑒 (5.5.-2)
𝑁𝐻3 𝑎𝑞 + 𝐻+ − 𝑧𝑒𝑜𝑙𝑖𝑡𝑒 ⇄ 𝑁𝐻4+ − 𝑧𝑒𝑜𝑙𝑖𝑡𝑒 (5.5.-3)
21
Some commonly used natural zeolites for removing ammonia are clinoptilolite and mordenite. These
zeolites will be compared by their ability to adsorb ammonia, that is, their ammonia uptake (AmU)
ability measured as [mg ammonia/g zeolite]. [70]
5.5.1. Mordenite
The influence of the cat ion on mordenites AmU is very diverse. Choosing the best cat ion is therefore
of great importance. As seen in table 12 the ion exchange level, defined as 𝐴𝑚𝑈
𝑡𝑒𝑜𝑟𝑒𝑡𝑖𝑐𝑎𝑙 𝐴𝑚𝑈, varies
from about 10 % up to 80 % depending on what cation the zeolite is prepared with. [70]
Table 12 Showing the influence of cation on mordenites AmU. [70]
Zeolite AmU [mmol/g] AmU [mg/g] Ion exchange level [%]
H-mordenite 0.15 2.56 9.8
K-mordenite 1.12 19.09 73.2
Na-mordenite 1.21 20.61 79.1
The temperature dependency has also been investigated and found not to be of importance in the
temperature range of 5-60 oC. [70] When using mordenite as zeolite 𝑁𝑎+ should be the cation of
choice giving an AmU of 20.6 mg/g.
5.5.2. Clinoptilolite
Clinoptilolite is another natural zeolite and its composition and thereby also its ammonia uptake
ability will depend on where it is mined. For clinoptilolite mined in the Transcarpathian region,
Ukraine, a study has shown that the order of ammonium ion exchange selectivity for cat ions are
𝑁𝑎+ > 𝐶𝑎2+ > 𝐾+. It was also shown that AmU for the Transcarpathian clinoptilolite varied in the
interval 13.56-21.52 mg/g depending on the flowrate of the solution (initial concentration 100 ppm).
[71]
6. Discussion of the literature part As has been investigated in this study contaminants will be formed and hence has to be removed.
The contaminants in question are ammonia, carbon dioxide and methanol. Possible ways to remove
these contaminants are steam and air stripping, ion exchange, deaeration and different biological
treatments. Biological treatment can be used to reduce the ammonia and methanol level. This
treatment method is very dependent on the temperature and the residence time is long which leads
to a need of very large cisterns. Another disadvantage is the long startup time for the process. In
deaeration 75 – 95 % of the carbon dioxide can be removed. These two separation methods are not
considered good enough and will therefore not be further investigated. Air stripping is not capable to
remove methanol and is therefore rejected. Steam stripping and ion exchange on the other hand has
been found to be applicable for this system.
Reports state that steam stripping reduces the methanol content with about 98 %. These results
have been accomplished for starting concentrations at about 5000 mg/kg. In this case the start
concentration of the condensate will be between 500 – 1000 mg/kg and therefore we assume that
98 % removal of methanol also can be applied.
22
Ion exchange can be used to remove the last trace concentrations of ammonia after the steam
stripper. The removal is depending on the amount and kind of zeolite and removal of about 20 mg/g
zeolite can be achieved.
Simulation part 7. Result and discussion
7.1. Shift simulation To investigate the gas composition in the effluent gas from the shift reactor and especially the
amount of by-products a simulation has to be set up. The starting values of the simulation is the
given gas composition of the reformer effluent. Since no data of the amount of ammonia is given it
has to be calculated before the shift reactor can be simulated.
Because the amount of ammonia always is at equilibrium [34] at the temperatures of the reformer,
no kinetic model of the reformer has to be set up. Instead an equilibrium equation for given
temperature out from the reformer is used. The equilibrium equation given by Gillespie and Beattie
[35] has been used to calculate the equilibrium composition of ammonia as described below:
𝐾𝑝 =𝑝𝑁𝐻3
𝑝𝐻21.5 ∙𝑝𝑁2
0.5 (7.-eq1)
Where Kp is calculated by equation (3.2.-eq3) and (3.2.-eq4).
After adding the calculated amount of ammonia and decreasing the amount of nitrogen and
hydrogen with equivalent amounts, this composition is the start concentration in the shift
simulation.
The shift reactor is simulated as an integral reactor using Skrzypek et al.’s equations described in
section 3.3. The differential equation that was used is seen below.
𝑑𝐶𝑀𝑒𝑂𝐻
𝑑𝑡= 𝑟1 (7.-eq2)
𝑑𝐶𝐻2𝑂
𝑑𝑡= 𝑟1 − 𝑟2 (7.-eq3)
𝑑𝐶𝐶𝑂2
𝑑𝑡= −𝑟1 + 𝑟2 (7.-eq4)
𝑑𝐶𝐻2
𝑑𝑡= −3 ∙ 𝑟1 + 𝑟2 (7.-eq5)
𝑑𝐶𝐶𝑂
𝑑𝑡= −𝑟2 (7.-eq6)
The reason why Skrzypek et al.’s equations were chosen among all those given in section 3.3 is
because most of them were missing values of the constants and no data were given to tune these
constants. The only thing missing in Skrzypek et al.’s equations is the value of the equilibrium
constants for both reactions and these have been taken from the program HSC [72] for given
pressure and as function of temperature.
23
The temperature is increasing along the reactor, to handle this an assumption has to be made about
the temperature profile. In the start the reaction rate is fast because it is far from equilibrium but
slow since the temperature is low and in the end it is the opposite. Therefore the assumption that
the temperature profile is linearly rising along the reactor is made. This is not entirely correct but it is
a good enough assumption and besides the focus of this work is not to simulate the shift reactor in
detail but to investigate wastewater purification.
Another problem is that the catalyst that Skrzypek et al.’s equations are constructed for is a catalyst
from 1985 and the ones existing today are probably more effective. To compensate for this more
mass of catalyst has been used in the simulation than in the real reactor. The used mass of catalyst in
the simulation is 3.61 times higher than in the real shift reactor. This value is used because then the
result from the simulation in regard to concentration of carbon monoxide, carbon dioxide, hydrogen
and water is almost the same as the data given from Haldor Topsoe A/S.
The result from the simulation is given in table 13 below.
Table 13 Mole composition and flow from shift reactor
Composition, dry basis Molar flow [kmol/h]
Methane 5.36 mole % 645
Carbon monoxide 3.61 mole% 435
Carbon dioxide 16.93 mole % 2039
Hydrogen 72.6 mole % 8746
Nitrogen 1.36 mole % 164
Ammonia 308 mole ppm 3.71
Methanol 831 mole ppm 10.0
Water - 1735
7.2. Stripping To model the plant given from Haldor Topsoe A/S in Aspen Plus some simplifications were made, the
cooling of the shift effluent and the heating of the condensate are modelled to take place in a single
heat exchanger. To model the heat exchanger the unit Heater in Aspen is used. The condenser is
modelled as a Flash2 and for the stripper a Radfrac is used. In the simulations the trays of the
stripper are considered ideal, the conversion to real trays is done when the cost of the vessel is
calculated. The layout is shown in figure 12. The flows and amounts of the different compounds can
be found in appendix A; stream 6545 is not included since it varies between the different treatment
configurations. Mass balances over the units show great compliance for all of them. The largest
deviation is for the stripper and is maximum 0.7 %. A more detailed description of the simulation
setup can be found in appendix D.
24
Figure 12 Aspen flow sheet with labelling of all the streams and units.
Three different stripper configurations were tested. At first a conventional single column stripper
with a fresh steam feed. In the second configuration no steam was fed to the column, instead a
reboiler was added. Then the effect of pH was tested by adding sodium hydroxide to a steam stripper
configuration and thus raising the pH. The purification efficiency is a measure of how much of the
impurities that are removed. It is calculated as.
𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑐𝑦 = 𝑎𝑚𝑜𝑢𝑛𝑡 𝑖𝑛−𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑢𝑡
𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑢𝑡 (7.4.-eq1)
The efficiencies investigated are 95, 99 and 99.9 %. The corresponding concentration levels are
shown in table 16.
Table 14 Effluent amount of the impurities for different purification efficiencies
Efficiency level CO2 [ppmw] NH3 [ppmw] MeOH [ppmw]
95 % 455 105 407
99 % 91 20.9 81.4
99.9 % 9.1 2.09 8.14
The input to the strippers is shown in table 17.
25
Table 15 Influent to the stripper section.
Stream 6530
Total flow [kg/h] 31900
Temperature [°C] 208
Pressure [bar] 32.8
Concentration CO2 [ppmw] 9100
Concentration NH3 [ppmw] 2090
Concentration MeOH [ppmw] 8140
An example of the results from Aspen of the system before the strippers can be found in appendix A.
The results used to create the different graphs and calculations can be found in appendix B.
7.2.1. Steam stripper
To simulate a steam stripper a multi-stage column with the influent process condensate on the top
tray and the influent steam on the bottom tray was set up. Efficiency of the steam stripper is mainly
depending on the number of trays and steam to feed ratio used. Also the pH will affect strongly on
the purity level reached for a certain number of trays.
To evaluate the purity dependency of number of trays installed in the column simulations were made
for a range of 2 – 20 ideal trays. The feed of steam was kept at 12 349 kg/h, which is the steam flow
given from the Haldor Topsoe A/S plant, in these simulations resulting in a steam to feed ratio of
about 0.38. The concentration of contaminants in the effluent process condensate from the stripper
is shown in figure 13 - 15 below.
Figure 13 Simulated concentration (logarithmic) of carbon dioxide in the effluent condensate for a range of installed trays
1.00E-35
1.00E-30
1.00E-25
1.00E-20
1.00E-15
1.00E-10
1.00E-05
1.00E+00
2 4 6 8 10 12 14 16 18 20
Co
nce
ntr
atio
n [
pp
mw
]
Installed trays
Carbon dioxide
CO2
26
Figure 14 Simulated concentration of ammonia in the effluent condensate for a range of installed trays
Figure 15 Simulated concentration of methanol in the effluent condensate for a range of installed trays
As can be seen in the figures above reduction in concentration of contaminants is achieved when
installing more trays but the reduction per tray is decreasing with increasing number of trays. For
ammonia 95 % purification efficiency would need 5 ideal trays while 99 % would need 9 ideal trays
and 99.9 % would need 14 ideal trays. The methanol is slightly harder to clean and needs 6 ideal trays
for 95 % purification efficiency, 9 ideal trays for 99 % and 15 ideal trays for 99.9 %. As seen in figure
13 the simulation shows that only trace concentrations of carbon dioxide will be found in this range
of installed trays.
To evaluate the purity dependency of steam to feed ratio in the column simulations were made for
ratios ranging from 0.2 to 0.38. The concentration of contaminants in the effluent process
condensate from the stripper is shown in figure 16 – 18 below.
0.0050.00
100.00150.00200.00250.00300.00350.00400.00450.00500.00
2 4 6 8 10 12 14 16 18 20
Co
nce
ntr
atio
n [
pp
mw
]
Installed trays
Ammonia
NH3
0.00200.00400.00600.00800.00
1000.001200.001400.001600.001800.002000.00
2 4 6 8 10 12 14 16 18 20
Co
nce
ntr
atio
n [
pp
mw
]
Installed trays
Methanol
CH3OH
27
Figure 16 Simulated concentrations (logarithmic) of carbon dioxide in the effluent condensate for a range of steam to feed ratios
Figure 17 Simulated concentration of ammonia in the effluent condensate for a range of steam to feed ratios
1.0E-34
1.0E-30
1.0E-26
1.0E-22
1.0E-18
1.0E-14
1.0E-10
1.0E-06
1.0E-02
0.2 0.25 0.3 0.34 0.3824
Co
nce
ntr
atio
n [
pp
mw
]
Steam to feed ratio
Carbon dioxide
5 trays
10 trays
15 trays
20 trays
0
100
200
300
400
500
600
700
0.2 0.25 0.3 0.34 0.3824
Co
nce
ntr
atio
n [
pp
mw
]
Steam to feed ratio
Ammonia
5 trays
10 trays
15 trays
20 trays
28
Figure 18 Simulated concentration of methanol in the effluent condensate for a range of steam to feed ratios
The need for steam decreases when increasing the number of installed trays as can be seen in the
figures above. For ammonia the steam to feed ratio can be reduced from 0.38 to about 0.25 by
increasing installed trays from 5 to 20 for an effluent concentration of about 100 ppmw (about 95 %
purification demand). This would decrease the steam usage from 12349 kg/h to 8073 kg/h, which is a
decrease of about 35 %. Methanol however is still harder to strip off than ammonia. For an effluent
concentration of about 420 ppmw (about 95 % purification demand) the steam to feed ratio can be
reduced from 0.38 to about 0.28 by increasing installed trays from 5 to 20, which is a decrease of
about 27 %.
Carbon dioxide on the other hand shows a whole different behaviour, the steam to feed ratio seems
to be of low importance for few installed trays. Figure 16 shows that greater reduction in carbon
dioxide concentration is achieved by installing more trays rather than using more steam, but the
influence of steam to feed ratio increases with more trays.
7.2.2. Stripper with reboiler
Instead of adding steam to the column a kettle type reboiler is placed in the bottom of it. The
reboiler contributes energy to the column. For the simulation the number of trays and bottom to
feed ratio have been varied. The bottom to feed ratio (reboil ratio) is defined as the ratio between
the liquid molar flow rate in the bottom to the molar feed rate. It can be calculated with the
following equation:
𝑅𝑎𝑡𝑖𝑜 =𝐹𝐵𝑜𝑡𝑡𝑜𝑚
𝐹𝐹𝑒𝑒𝑑 (7.2.2.-eq1)
In the first simulation different numbers of trays and different bottom to feed ratio were investigated
to see how the system reacts to the different settings. The result can be seen in figure 19 – 21.
0
500
1000
1500
2000
2500
3000
0.2 0.25 0.3 0.34 0.3824
Co
nce
ntr
atio
n [
pp
mw
]
Steam to feed ratio
Methanol
5 trays
10 trays
15 trays
20 trays
29
Figure 19 The logarithmic carbon dioxide content in the effluent condensate
In figure 19 the amount of carbon dioxide in the condensate is shown. It can be seen that at high
reboil ratio the carbon dioxide content varies. If the reboil ratio is 0.9 the amount of carbon dioxide is
below 1 ppmw. It will be very easy to remove the carbon dioxide and a stripper with 5 trays and a
reboil ratio of 0.9 will give a carbon dioxide content of 0.06 ppmw and the purification level will be
very close to 100 %.
Figure 20 The ammonia content in the effluent condensate
As can be seen in figure 20 the ammonia content is very high in the condensate at bottoms to feed
ratio at 0.95 and there are a very small differences in the ammonia content between the different
numbers of trays. At a reboil ratio of 0.7 and 20 trays the ammonia content in the condensate is 3.77
ppmw which gives a purification efficiency of 99.9 %. If the number of trays is 15 the ammonia
1.00E-33
1.00E-30
1.00E-27
1.00E-24
1.00E-21
1.00E-18
1.00E-15
1.00E-12
1.00E-09
1.00E-06
1.00E-03
1.00E+00
Co
nce
ntr
atio
n [
pp
mw
]
Reboil ratio
Carbon dioxide
5 trays
10 trays
15 trays
20 trays
0.00E+00
2.00E+02
4.00E+02
6.00E+02
8.00E+02
1.00E+03
1.20E+03
1.40E+03
1.60E+03
1.80E+03
2.00E+03
Co
nce
ntr
aio
n [
pp
mw
]
Reboil ratio
Ammonia5 trays
10 trays
15 trays
20 trays
30
content will be 11.9 ppmw and the purification efficiency will be 99 % and if the number of trays is 10
the content will be 49.9 ppmw and the purification efficiency will be 95 %.
Figure 21 Methanol content in the effluent condensate
Figure 21 shows how the methanol content in the condensate varies at different bottoms to feed
ratio and different numbers of trays. If the reboil ratio is high the amount of methanol in the
condensate will be high. To reach a low level it can be seen in the figure 21 that the column must
have 10 – 20 trays and the boilup ratio must be around 0.7. If the number of trays is 10 the
purification efficiency will be 95 %. If the numbers of trays increases to 20 the purification efficiency
will be 99.9 %.
As can be seen in the figures the best result will be accomplished at high numbers of trays and low
reboil ratio. It looks like a reboil ratio of 0.7 will give good purification efficiency for a column with 10
– 20 trays. In the following three figures the tray dependency is investigated.
Figure 22 Logarithmic carbon dioxide content in the condensate at boilup ratio of 0.7 and different numbers of trays
0.00E+00
1.00E+03
2.00E+03
3.00E+03
4.00E+03
5.00E+03
6.00E+03
7.00E+03
8.00E+03
Co
nce
ntr
aio
n [
pp
mw
]
Reboil ratio
Methanol5 trays
10 trays
15 trays
20 trays
1.00E-33
1.00E-30
1.00E-27
1.00E-24
1.00E-21
1.00E-18
1.00E-15
1.00E-12
1.00E-09
1.00E-06
1.00E-03
1.00E+00
5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20
Co
nce
ntr
atio
n [
pp
mw
]
Number of trays
Carbon dioxide
31
As can be seen in figure 22 the carbon dioxide content is 1.35·10-4 ppmw at bottoms to feed ratio of
0.7 and 5 trays which gives a purification efficiency of 100 %.
Figure 23 The ammonia content in the condensate at a reboil ratio of 0.7 and different numbers of trays
In figure 23 the bottoms to feed ratio is set to 0.7 and the numbers of trays has been varied. As
mentioned before a high number of trays will give the best result. 20 trays will give a level of 2.89
ppmw and purification efficiency of 99.9 %, 15 trays will give a purification efficiency of 99 % and 8
trays will give a purification of 95 %.
Figure 24 Methanol content in the condensate at reboil ratio of 0.7 and different numbers of trays
In figure 24 methanol is shown at reboil ratio of 0.7. A column of 20 trays would give an amount of
19.9 ppmw methanol in the condensate and a purification efficiency of 99.9 %, 15 trays will give a
purification of 99 % at a level of 70 ppmw and 8 trays will give an purification of 95 %.
As can be seen in the past figures the best result can be accomplished with a column of 20 trays. To
see the impact of the reboil ratio a simulation of different reboil ratios has been done. The result is
shown in the following figures.
0.00E+00
5.00E+01
1.00E+02
1.50E+02
2.00E+02
2.50E+02
5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20
Co
nce
ntr
atio
n [
pp
mw
]
Number of trays
Ammonia
0.00E+00
2.00E+02
4.00E+02
6.00E+02
8.00E+02
1.00E+03
1.20E+03
5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20
Co
nce
ntr
atio
n [
pp
mw
]
Number of trays
Methanol
32
In a column of 20 trays the carbon dioxide content will be very low and the purification efficiency will
be 100 %. To get the optimal reboil ratio different reboil ratios has been simulated for a column with
20 trays.
Figure 25 Ammonia content at 20 trays and different reboil ratios
Figure 25 shows the ammonia content in the condensate. At a reboil ratio of 0.69 the ammonia
content in the outlet is 1.85 ppmw and the purification efficiency reaches a level of 99.9 %. A reboil
ratio of 0.73 will give an ammonia content of 12.6 ppmw and a purification of 99 %, if the reboil ratio
is set to 0.77 the ammonia content will be 95.6 ppmw and the purification will be 95 %.
Figure 26 Methanol content at 20 trays and different boilup ratios
The methanol content decreases with the reboil ratio which is shown in figure 26. At a reboil ratio of
0.69 the methanol amount is 11.7 ppmw which gives a purification efficiency of 99.9 %. At a reboil
ratio of 0.73 the amount of methanol will be 95.9 ppmw and 99 % purification, a reboil ratio of 0.75
will give an amount of 271 ppmw and a purification level of 95 %.
0.00E+00
5.00E+01
1.00E+02
1.50E+02
2.00E+02
2.50E+02
3.00E+02
3.50E+02
0.8 0.79 0.78 0.77 0.76 0.75 0.74 0.73 0.72 0.71 0.7 0.69 0.68 0.67
Co
nce
ntr
atio
n [
pp
mw
]
Reboil ratio
Ammonia
0.00E+00
2.00E+02
4.00E+02
6.00E+02
8.00E+02
1.00E+03
1.20E+03
1.40E+03
1.60E+03
1.80E+03
2.00E+03
0.8 0.79 0.78 0.77 0.76 0.75 0.74 0.73 0.72 0.71 0.7 0.69 0.68 0.67
Co
nce
ntr
atio
n [
pp
mw
]
Reboil ratio
Methanol
33
It can be seen in the figures that when the reboil ratio decreases the amount of impurities decreases.
The drawback is that more power is needed which can be seen in figure 27. An increase in power
demand will increase the costs. Another drawback is that the condensate flow will decrease.
Figure 27 Power demand at different reboil ratios
7.2.3. Steam stripper with pH adjusting
The third set-up for the stripper system was one with two subsequent columns, figure 28. The first
column was designed with five and the second with ten ideal trays. In order to raise pH, in
accordance with the literature cited, an extra stream containing sodium hydroxide, caustic, was
added. The stream was fed to the first tray, the upper one, in the second column. Aspen was then set
to calculate the amount of caustic needed to obtain the desired pH value at the first tray.
The reasoning behind this set-up is to drive off the carbon dioxide, favoured by a lower pH as argued
previously, in the first column and then the ammonia, favoured by higher pH, in the second one. The
pH in the system before the stripper is just above 7 which is well below the, in the literature,
recommended pH which is close to 11.
One thing to keep in mind when adjusting the pH is the buffering capacity of the system. Both carbon
dioxide and ammonia have buffering capacities which works in opposite directions though, see
equations 5.3.-2 – 5.3.-5. Before the pH can be adjusted the buffering capacity has to be neutralised,
this is actually the whole purpose of raising pH, shifting the equilibrium of ammonia and ammonium
towards ammonia. In this set-up the concentration of carbon dioxide entering stripper 2 is much
lower than the concentration of ammonia. The concentration of carbon dioxide is about 0.25 % of
the ammonia concentration at the most. Therefore the buffering effect of carbon dioxide is
neglected. The amount of ammonia entering stripper 2 varied between 1.3 and 3.3 kmol/h. To
neutralise the ammonia, assuming all ammonia is in the form of ammonium, equal amounts of
caustic is needed. When the buffering capacity is neutralised the amount of caustic needed to raise
pH is negligible giving a demand of caustic between 50 and 130 kg/h, which are about 0.15 – 0.4 % of
the total flow.
0
1
2
3
4
5
6
7
0.8 0.79 0.78 0.77 0.76 0.75 0.74 0.73 0.72 0.71 0.7 0.69 0.68 0.67
Po
we
r d
em
and
[M
W]
Reboil ratio
Energy
34
Figure 28 Aspen flow sheet with labelling of all the streams and units for the double stripper configuration.
The simulations however gave a totally different result. In order to raise pH to 11 the simulations
indicate a need of about 8 MT/h, 25 % of the feed, which of course is not correct. Why the
simulations did not give a more accurate answer we cannot explain. For the sake of being able to
retrieve any results we have assumed that the results, for everything except the amount of caustic,
are valid. The results are of the same magnitude as those from the tests without caustic which
supports the assumption. The amount of caustic needed is instead assumed to the one from the
rough calculation above.
The most important design parameter is the amount of fresh steam added to the strippers and the
distribution of it between the two columns. Especially the distribution greatly affects the outcome. A
large steam flow in column one gives a greater reduction of carbon dioxide whereas a larger flow in
the second column gives a larger reduction of ammonia and methanol. This is illustrated in the
following three figures, 29 – 31. In the figures the amount of steam fed to column 1 is reduced in run
1 to 4 and the amount fed to column 2 is increased thus giving a constant total steam flow to the
system, table 16.
Table 16 Steam distribution between column 1 and 2 in figure 29 – 31
Run Steam ratio [S:F 1/S:F 2] Total steam [kg/h]
1 1.0 13000
2 0.75 13000
3 0.54 13000
4 0.38 13000
Temperature (C)
Pressure (bar)
Mass Flow Rate (kg/hr)
Vapor Fraction
HEX1
FLASH
HEX2
STRIPP1
329
23.1
168718
1.00
2300
40
23.1
168718
0.87
2360
40
23.1
136782
1.003200
40
23.1
31936
0.00
6500
208
32.8
31936
0.03
6530236
32.8
32616
0.00
6545
209
32.8
1904
1.006340
253
41.5
2585
1.00STEAM
STRIPP2
253
41.5
7415
1.00STEAM2
249
32.8
11117
1.00VAPOUT
250
32.8
37099
0.00
LIQOUT
238
32.8
8185
0.00CAUSTIC
35
Figure 29 The effect of different steam distributions between column 1 and 2 for carbon dioxide removal. In run no. 1 the distribution is even, in run 2 the ratio between column 1 and 2 is 0.75, in run 3 the ratio is 0.54 and in run 4 the ratio is 0.38. The total amount of fresh steam is 13000 kg/h. The concentration is the concentration in the bottom effluent of the second column.
Figure 30 The effect of different steam distributions between column 1 and 2 for ammonia removal. In run no. 1 the distribution is even, in run 2 the ratio between column 1 and 2 is 0.75, in run 3 the ratio is 0.54 and in run 4 the ratio is 0.38. The total amount of fresh steam is 13000 kg/h. The concentration is the concentration in the bottom effluent of the second column. The line for pH 11 is located behind the one for pH 10.5.
00.020.040.060.08
0.10.120.140.160.18
0.2
1 2 3 4
Co
nce
ntr
atio
n [
pp
mw
]
Run no.
Carbon dioxide
pH 9
pH 10
pH 10.5
pH 11
0102030405060708090
100110120130
1 2 3 4
Co
nce
ntr
atio
n [
pp
mw
]
Run no.
Ammonia
pH 9
pH 10
pH 10.5
pH 11
36
Figure 31 The effect of different steam distributions between column 1 and 2 for methanol removal. In run no. 1 the distribution is even, in run 2 the ratio between column 1 and 2 is 0.75, in run 3 the ratio is 0.54 and in run 4 the ratio is 0.38. The total amount of fresh steam is 13000 kg/h. The concentration is the concentration in the bottom effluent of the second column.
The effect of changing the distribution is largest for carbon dioxide on a percent basis. The levels are
however much lower, up to 1000 times, compared to the other two. One can argue if the change is
of any real significance for carbon dioxide because of the low concentrations, 10-5 – 0.2 ppmw. For
the other two the effect is of greater importance, especially at lower pH and not as much at the
higher ones. In absolute amounts the distribution is of greatest importance for the methanol
removal.
It is also evident that pH affects the results although in different ways. Carbon dioxide is affected,
negatively, when pH is above 9. The effect is increased when the steam feed is reduced. The effect is
basically the same for all pH above 9. Again one can argue if differences at concentration levels as
low as this is of any importance. Both ammonia and methanol is favoured by a higher pH. There is a
drastic improvement when pH is changed from 10 to 10.5. The fact that ammonia is greatly
influenced by pH is expected. Why methanol on the other hand is influenced is more uncertain. As
far as we know methanol should not be affected by pH since it has no ionic forms and does not bind
in to any complexes, which is not included in our Aspen model anyway. One explanation could be
that methanol removal is hampered by the presence of ammonia which is more volatile, when the
ammonia is removed so is the methanol.
Next it was investigated what happens if the total amount of steam is reduced. A number of different
total steam flows were tested. Two different distributions between the two columns were also
tested, the results is shown in figures 32 to 35. The two distributions tested were either a steam to
feed ratio to column 1 of 0.080 or 0.12, the steam to feed ratio in column 2 were adjusted to give the
right total amount of steam fed to the system. As seen in figure 29 - 31 it is difficult to get an
effective cleaning with pH 9 and 10 even at a higher total steam feed. Therefore no further tests at
those pH´s were done.
050
100150200250300350400450500550600650
1 2 3 4
Co
nce
ntr
atio
n [
pp
mw
]
Run no.
Methanol
pH 9
pH 10
pH 10.5
pH 11
37
Figure 32 Bottom concentration of carbon dioxide from the second stripper for different pH and steam distributions between the two columns. The steam to feed ratio in column 1 is either 0.08 (upper set of values) or 0.12 and the amount steam to column 2 is adjusted after that given the right total amount of steam.
Figure 33 Bottom concentration of ammonia from the second stripper for different pH and steam distributions between the two columns. The steam to feed ratio in column 1 is either 0.08 (lower set of values) or 0.12 and the amount steam to column 2 is adjusted after that given the right total amount of steam.
0
0.5
1
1.5
2
2.5
3
3.5
4
4.5
9500 10000 10500 11000 11500 12000 12500
Co
nce
ntr
atio
n [
pp
mw
]
Total amount of steam [kg/h]
Carbon dioxide
pH 10.5
pH 11
0
2
4
6
8
10
12
14
16
18
9500 10000 10500 11000 11500 12000 12500
Co
nce
ntr
atio
n [
pp
mw
]
Total amount of steam [kg/h]
Ammonia
pH 10.5
pH 11
38
Figure 34 Bottom concentration of methanol from the second stripper for different pH and steam distributions between the two columns. The steam to feed ratio in column 1 is either 0.08 (lower set of values) or 0.12 and the amount steam to column 2 is adjusted after that given the right total amount of steam.
Once again it is evident how important the distribution of the steam between the two columns is.
When the amount of steam fed to the first column is reduced by changing the steam to feed ratio,
the bottom effluent concentration is increased for carbon dioxide and decreased for ammonia and
methanol. Figure 32 confirms the results from figure 29 that pH does not affect the carbon dioxide
removal to any greater extent, at least not for the higher steam flow in column 1. Once again the
removal of ammonia and methanol is favoured by a higher pH. A higher pH also seems to dampen
the effects of changing the steam distribution between the two columns. At pH 11 there is no
significant difference between the two distributions with regard to ammonia. For a feed of 11 or 12
tonnes of steam per hour the difference is not that big in absolute numbers for methanol either. This
indicates that the distribution could be chosen with regard to carbon dioxide if the steam feed is at
least 11 tonnes per hour and with a pH of at least 11.
The purification efficiency is above 95 % for ammonia in all tests but one, run 1 at pH 9 in figure 30. It
is also above 99 % if pH is 10.5 or higher. Depending on the distribution of steam between the two
columns an efficiency of 99.9 % can be met at pH 10.5 and a total amount of steam of 11000 kg/h.
For methanol an efficiency of 99.9 % can be reached as long as the pH is adjusted to 11. 99 %
efficiency can be met if the pH is 10.5 and the total amount of steam is 12000 kg/h. An efficiency of
95 % can be met at a pH of 9. Notable is that the efficiency is dependent on the distribution of steam
between the two columns at a given total steam flow; the efficiency can vary up to five percentage
points. Thus, to meet the efficiency levels above the distribution has to be properly adjusted in some
of the cases. For carbon dioxide the purification efficiency is above 99.9 % in all of these runs.
7.3. Ion exchanger L. R. Weatherley and N. D. Miladinovic have developed a Langmuir expression for the uptake of
ammonia in clinoptilolite, as shown in (7.3.-eq1) below. [75, 76]
𝑄𝑒 =9.835∙𝐶𝑒
1+1.493∙𝐶𝑒 (7.3.-eq1)
050
100150200250300350400450500
9500 10000 10500 11000 11500 12000 12500
Co
nce
ntr
atio
n [
pp
mw
]
Total amount of steam [kg/h]
Methanol
pH 10.5
pH 11
39
For ammonia concentration of 𝐶𝑒 = 10 𝑚𝑔/𝑙 the capacity of the clinoptilolite is
𝑄𝑒 = 6.17 𝑚𝑔𝑁𝐻4+−𝑁/𝑔𝐶𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 .
Water flow from bottom of the stripper is 30 ton/h, which for a linear velocity which is approximated
to 4 cm/min in the bed gives a total area of 12.5 m2.
𝐴𝑝𝑎𝑐𝑘𝑖𝑛𝑔 =𝐹
𝑣 (7.3.-eq2)
This gives, for a run time of 200 hours before regeneration, a needed mass of packing material of
10.8 MT, if it assumes to be 90 % saturated when regeneration starts.
𝑚𝐶𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 =𝐶𝑒 ∙𝑉 ∙𝑡
𝑄𝑒 ∙0.9 (7.3.-eq3)
This gives a packing height of 86.4 cm.
𝑝𝑎𝑐𝑘𝑖𝑛𝑔 =𝑚𝐶𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒
𝜌𝐶𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 ∙𝐴𝑝𝑎𝑐𝑘𝑖𝑛𝑔 (7.3.-eq4)
Where the density of clinoptilolite is 1000 kg/m3. [77]
If the diameter of the vessels is 2 meters, then the number of required vessels is four. But the ion
exchanging material must, when becoming saturated, be regenerated so the number of vessels has
to be five instead, thus allowing one vessel to be regenerated whilst the other ones are operating.
𝑁𝑢𝑚𝑏𝑒𝑟 𝑜𝑓 𝑣𝑒𝑠𝑠𝑒𝑙𝑠 =𝑚𝐶𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒
𝜌𝐶𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 ∙𝐴𝑣𝑒𝑠𝑠𝑒𝑙 ∙𝑝𝑎𝑐𝑘𝑖𝑛𝑔+ 1 (7.3.-eq5)
The total amount of clinoptilolite, including the packing in the extra vessel, becomes 13.5 MT.
𝑚𝑐𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 𝑡𝑜𝑡𝑎𝑙 = 𝑚𝑐𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 ∙ 1.25 (7.3.-eq6)
7.4. Verification of simulation results To verify the equilibrium composition obtained in Aspen Plus the program HSC [72] were used. HSC
has an extensive thermo chemical database and calculates equilibrium compositions for various
systems including reactions. In this project the equilibrium module were utilized. This module is
capable of calculating the equilibrium composition for heterogeneous multi-component system, such
as the one at hand. The calculations were done with the GIBBS solver which minimizes the Gibbs free
energy for the system.
Two different parts of the flow sheet were looked upon, the cooling of the shift-effluent before the
condenser and the heating of the condenser-effluent before entering the stripper. The program
handles the pure species in either gaseous or aqueous form; it is also capable of handling ionic forms.
One drawback was the fact that the program did not have any data for aqueous ammonia. This will of
course result in some differences in the results between the two programs. The amounts, the same
as in Aspen, were typed in and then the pressure and temperature were varied. The results are
shown in table 17 and 18 for the cooling and the heating respectively. Notable is that there is no way
of knowing how much of a species that is present in which form in Aspen, e.g. how much carbon
dioxide is in aqueous and gaseous form respectively.
40
Table 17 Cooling of shift-effluent before entering the condenser
HSC and Aspen
[kmol/h]
HSC
[kmol/h]
Aspen
[kmol/h]
Stream 2300 2360 2360
Temperature (°C) 329 40 40
Pressure (bar) 23.1 23.1 23.1
CO2(g) 2070 2050 2063*
H2O(g) 1750 6,41 -
NH3(g) 9.82 1,53E-03 0.0810*
H2O 0 1730 1741*
CO2(aq) 0 17,3 -
NH4+(aq) 0 9,81 9.73
HCO3-(aq) 0 9,81 9.72
CO32-(aq) 0 6,67E-04 7.03E-03
OH-(aq) 0 1,03E-06 5.73E-06
H+(aq) 0 2,83E-05 8.25E-06
*Total amount of the species, i.e. in both gaseous and aqueous phase.
Table 18 Heating of condenser-effluent before entering the stripper
HSC and Aspen
[kmol/h]
HSC
[kmol/h]
Aspen
[kmol/h]
Stream 6500 6530 6530
Temperature (°C) 40 208 208
Pressure (bar) 23.1 32.8 32.8
CO2(g) 0 7.57 11.8**
H2O(g) 0 14.7 -
NH3(g) 5.48E-02* 7.17 9.19**
H2O 1699 1690 1708**
CO2(aq) 2.64 2.18 -
NH4+(aq) 9.73 2.61 0.596
HCO3-(aq) 9.72 2.61 0.594
CO32-(aq) 7.03E-03 1.48E-03 1.06E-04
OH-(aq) 5.73E-06 3.32E-03 1.61E-03
H+(aq) 8.25E-06 1.45E-06 4,11E-06
* In Aspen the ammonia was aqueous form
**Total amount of the species, i.e. in both gaseous and aqueous phase.
It seems that the larger the amount of the species the greater the congruity. For the lower
concentration of species the difference is larger. The concentration of the carbonate ion for example
differs by a factor ten for the heating in table 18. This difference could depend on different
calculation methods for the different programs. HSC is strictly based equilibrium composition, the
system is assumed ideal, whereas Aspen also takes non-ideality into account. Aspen also utilises
different activity factor models. The values is still of the same magnitude and does not display any
exceptional differences, which to some degree confirms our calculations.
41
An analysis was also conducted on the importance of pressure; could the carbon dioxide be removed
simply by decreasing the pressure? As figure 35 shows no complete reduction of aqueous carbon
dioxide or its different ions were possible without the water also entering the gaseous phase, figure
36. The pressure had to be reduced to a point where the temperature of the liquid was equal to the
boiling point of water. The same goes for ammonia, figure 37. Reducing the pressure could still be of
interest since it is possible to reduce the aqueous forms of carbon dioxide and ammonia respectively
to almost half of the starting amount without the water being vaporised.
Figure 35 The effect of pressure reduction on the amount of the different forms of carbon dioxide
Figure 36 The effect of pressure reduction on the amount of water in gaseous form
0.00
2.00
4.00
6.00
8.00
10.00
12.00
14.00
32
.8
31
.4
30
.0
28
.6
27
.2
25
.8
24
.4
23
.0
21
.6
20
.2
18
.8
17
.4
16
.0
14
.7
13
.3
11
.9
10
.5
Am
ou
nt
[km
ol]
Pressure [bar]
Carbon dioxide
CO2 (g)
CO2 (a)
HCO3-
CO3-2
0
200
400
600
800
1 000
1 200
1 400
1 600
1 800
32
.8
31
.4
30
.0
28
.6
27
.2
25
.8
24
.4
23
.0
21
.6
20
.2
18
.8
17
.4
16
.0
14
.7
13
.3
11
.9
10
.5
Am
ou
nt
[km
ol]
Pressure [bar]
Water
H2O (g)
42
Figure 37 The effect of pressure reduction on the amount of the different forms of ammonia
7.5. Discussion If we compare the column with a kettle reboiler to the steam stripper the important thing is the heat
duty. In the steam stripper the heat is supplied as steam whereas in the reboiler stripper it is added
as external energy. With the same energy added the steam stripper is more efficient. This partly
because the added amount of steam makes the condensate more dilute. The steam that is utilised is
produced using waste heat from the reformer. Since the heat has to be cooled off anyway using it to
create steam reduces the need of cooling water.
When comparing the two steam stripping alternatives the result is dependent on which level of
impurities is acceptable for the different compounds. If it is most crucial to remove the carbon
dioxide steam stripping without adding caustic is more efficient reaching levels as low as 7·10-34
ppmw. The level of ammonia is 0.14 ppmw and the level of methanol is 0.71 ppmw.
If on the other hand ammonia is the biggest concern adding caustic should be considered. Our
simulations shows that levels as low as 4·10-5 ppmw ammonia can be reached whilst still keeping a
low level of carbon dioxide, 0.09 ppmw. The steam flow is 12 000 kg/h. The level of methanol is
around 1 ppmw, methanol is considered the least harmful of the three compounds, at least with
respect to corrosion, studied in this report.
If the purification demand is not as strict, savings can be made both on the material by choosing a
smaller column and also on the steam consumption. To reach a purification efficiency of 99.9 % only
10 000 kg/h steam is needed if the pH is adjusted to 11, adding 130 kg/h caustic which is about 0.4 %
of the total flow. An even lower efficiency gives even bigger saving since a smaller column can be
used.
The simulations show that methanol is the impurity that is most difficult to remove. One explanation
is of course the high level, compared to ammonia, which enters the strippers. In order to get rid of
the methanol a large number of trays and a lot of steam is required. This render the need of a
subsequent ion-exchange step unnecessary since ammonia already is removed under those
conditions.
0.00
0.20
0.40
0.60
0.80
1.00
1.20
32
.8
31
.4
30
.0
28
.6
27
.2
25
.8
24
.4
23
.0
21
.6
20
.2
18
.8
17
.4
16
.0
14
.7
13
.3
11
.9
10
.5
Am
ou
nt
[km
ol]
Pressure [bar]
Ammonia
NH3 (g)
NH4+
43
The configurations of greatest interest are either a single steam stripper or two strippers with added
caustic. They can both reach purification efficiencies over 99.9 %, the difference lies in the total cost.
Since the purification requirements are not known several alternatives will be compared with respect
to economy. The levels listed will be 99 %, 99.9 % efficiency and also the optimal efficiency for the
two different configurations, one and two columns. Efficiencies below 99 % are not considered
industrially applicable.
8. Economy To calculate cost for the strippers the Ulrich method [78] was used. The calculations and results, as
well as the cost for caustic can be found in appendix C.
To calculate the total cost for four different purification levels, a total of seven stripper
configurations, the present value method were used. [81] The present price of the purification
alternatives are listed in table 19. The steam is assumed to be produced in an auxiliary boiler with the
efficiency 90 %, fuelled by natural gas. The price of the natural gas used is 698 USD/1000 m3, giving a
steam price of 356 USD/MT. The present values are calculated based on ten years depreciation
period with 15 % internal rate of return.
𝑆𝑛 = −𝐺 + 𝑎𝑖 ∙ 1 + 𝑋 −𝑖𝑁𝑖=1 (8.-eq1)
The strippers suggested in this report are to be integrated in an existing plant configuration and
replace existing condensate treatment equipment. It is therefore assumed that there is no need for
additional personnel in order to operate the new equipment. The same argumentation is applied
when it comes to utility equipment such as pumps and other operational costs.
Table 19 Total cost for different purification efficiencies and stripper configurations. The calculations are based on ten years depreciation period with 15 % internal rate of return.
Optimal
single column
Optimal
two columns
99 %
alt. 1
99 %
alt. 2
Configuration 1 column
34 trays
2 columns
9+17 trays
1 column
17 trays
1 column
25 trays
Conc. in stripper effluent
(CO2/NH3/MeOH) [ppmw]
(6.8·10-34/
0.14/0.71)
(0.087/
3.4·10-5/1.2)
(7.9·10-15/
9.6/49)
(7.2·10-24/
5.4/32)
Investment [USD] 2 111 000 1 954 000 1 160 000 1 596 000
Steam [ton/y] 103 700 100 800 103 700 91 100
Price caustic [USD/y] 0 705 600 0 0
Price steam [USD/y] 3 697 000 3 593 000 3 697 000 3 247 000
Total cost steam included [USD] 20 666 000 23 526 000 19 720 000 17 890 000
Total cost steam excluded [USD] 2 111 000 5 495 000 1 160 000 1 596 000
44
Table 19 cont.
99.9 %
alt. 1
99.9 %
alt. 2
99.9 %
alt. 3
Configuration 1 column 25 trays 1 column 34 trays 2 columns 9+17 trays
Conc. in stripper effluent
(CO2/NH3/MeOH) [ppmw] (2.5·10-25/1.1/5.9) (5·10-34/1.1/7.3) (3.8/1.1·10-4/3.4)
Investment [USD] 1 596 000 2 111 000 1 954 000
Steam [ton/y] 103 700 91 100 84 000
Price caustic [USD/y] 0 0 898 800
Price steam [USD/y] 3 697 000 3 247 000 2 993 000
Total cost steam included [USD] 20 150 000 18 410 000 21 490 000
Total cost steam excluded [USD] 1 596 000 2 111 000 6 465 000
9. Discussion The choice of configuration depends on the purification requirements for the different contaminants
and the price of the steam used.
If no ammonia can be accepted in the condensate an ion-exchanger unit is an alternative as a
complement to the stripper. The cost of an ion-exchanger unit is however large, comparable with a
34 tray column.
The accuracy of the simulations is of course dependent on the physical properties available. The
Redlich-Kwong (RK) equation of state used in the simulations was developed in 1949. Since then,
many modifications have been done. Some rough calculations show that the difference between RK
and Peng-Robinson is 2.5 % when the volume of pure water at 32.8 bar and 236 °C is calculated. The
difference is not considered significant. Still it would have been interesting to compare the results
with results from a more advanced equation of state. One thing worth discussing is if it is realistic to
get levels as low as simulated. For example concentrations as low as 10-34 ppmw for carbon dioxide is
reached.
For the shift simulations the largest source of error is the kinetic expressions used. The problem was
to get all the constants used in the more credible expressions found in the literature. A way around
this problem would be to calibrate the constants in those equations given in the literature, but no
data that could have been used for this purpose was given. However, the result from the simulations
did, even for the impurities, correspond well with the result found in literature. [34]
A potential error source is the human factor. Aspen is a somewhat complex program that demands a
lot from the user. We were not familiar with the program and all the different settings that are
available on fore hand. It is therefore hard for us to evaluate our model. Especially the results from
the stripper are hard to validate since there is no data available for comparison.
The economy calculations are based on various diagrams and rough approximation methods. This of
course means that the accuracy of the results is not that great.
One thing not answered in this report that is highly interesting is the relation between methanol
removal and pH. The results clearly show that pH greatly affects the removal efficiency; a higher pH
45
leads to more efficient methanol removal. In order to fully be able to practically implement the
results from this report a greater understanding of this phenomenon is of importance.
An interesting configuration is two columns of each ten ideal trays and caustic added to the second
column. This should allow very low levels of both carbon dioxide and ammonia to be reached. The
investment cost is not that much higher compared to a single 20 ideal trays column. This
configuration was not tested in this report. If further test are to be done we recommend testing this
configuration. Another thing that is worth investigating further is the amount of caustic needed for
the pH adjusting. The amounts in this report are based on rough calculations and therefore more
precise methods are desired in order to calculate the cost more exact.
10. Conclusion Our simulations show that it is possible to reach a very efficient purification of the condensate, above
99.9 %. Using steam stripping can give a condensate with very low levels of the impurities
investigated, levels as low as 10-34 ppmw is reached for carbon dioxide. The total cost for this
configuration is 21 million USD over a ten year period. Adding caustic is very effective to drive of
ammonia, levels of 10-5 ppmw is reached. The adding of caustic hampers the removal of carbon
dioxide, still very low levels of the carbon dioxide is reached, 0.087 ppmw, at a total cost of 24 million
USD over a ten year period. Methanol is mainly dependent on the amount steam added and is
proven more difficult to remove than the other two, levels of around 1 ppmw is reached for the
configurations above.
46
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a
Appendix
A. Example of Aspen results
Stream 2300 2360 3200 6500 6530
Mole Flow kmol/hr
H2O 1751.168 1751.168 44.25309 1706.915 1706.915
CH4 653.4868 653.4868 624.6434 28.84339 28.84339
CO 444.1944 444.1944 442.3946 1.799759 1.799759
CO2 2072.669 2072.669 2066.533 6.135878 6.135878
H2 8895.353 8895.353 8894.982 0.371812 0.371812
N2 168.246 168.246 168.2409 0.005108 0.005108
NH3 3.56926 3.56926 0.005128 3.564132 3.564132
CH3OH 9.815464 9.815464 1.762634 8.05283 8.05283
NH4+ 0 0 0 0 0
H3O+ 0 0 0 0 0
HCO3- 0 0 0 0 0
OH- 0 0 0 0 0
CO3-- 0 0 0 0 0
Mass Flow kg/hr
H2O 31547.78 31547.78 797.2318 30750.55 30750.55
CH4 10483.73 10483.73 10021 462.7275 462.7275
CO 12442.06 12442.06 12391.65 50.41197 50.41197
CO2 91217.75 91217.75 90947.71 270.0388 270.0388
H2 17931.96 17931.96 17931.22 0.749529 0.749529
N2 4713.155 4713.155 4713.012 0.143089 0.143089
NH3 60.78649 60.78649 0.087326 60.69916 60.69916
CH3OH 314.5087 314.5087 56.47859 258.0301 258.0301
NH4+ 0 0 0 0 0
H3O+ 0 0 0 0 0
HCO3- 0 0 0 0 0
OH- 0 0 0 0 0
CO3-- 0 0 0 0 0
Total Flow kmol/hr
13998.5 13998.5 12242.81 1755.688 1755.688
Total Flow kg/hr 168711.7 168711.7 136858.4 31853.35 31853.35
Total Flow cum/hr
30451.15 13889.53 14489.32 35.40998 98.45758
Temperature C 329 40 40 40 208
Pressure bar 23.11325 23.11325 22.1 22.1 32.81325
Vapor Frac 1 0.874348 1 0 0.030321
Liquid Frac 0 0.125652 0 1 0.969679
Solid Frac 0 0 0 0 0
Enthalpy kcal/mol
-20.509 -24.0411 -17.8598 -67.1315 -63.8581
Enthalpy kcal/kg -1701.69 -1994.76 -1597.66 -3700.14 -3519.72
b
Enthalpy Gcal/hr -287.1 -336.544 -218.657 -117.864 -112.116
Entropy cal/mol-K
0.004841 -8.53217 -4.18048 -38.1869 -29.6133
Entropy cal/gm-K 0.000402 -0.70794 -0.37397 -2.10478 -1.63222
Density kmol/cum
0.459704 1.007846 0.844954 49.58172 17.83192
Density kg/cum 5.540406 12.14668 9.445467 899.5586 323.5236
Average MW 12.05213 12.05213 11.17867 18.14295 18.14295
Liq Vol 60F cum/hr
687.4191 687.4191 654.1036 33.31548 33.31548
c
B. Results
Steam stripper
No. trays 2 3 4 5 6 7 8 9
CO2 out [ppmw] 4.156700369 0.095442561 0.002078299 4.26932E-05 8.28801E-07 1.20171E-08 1.08944E-10 9.39519E-13
NH3 out [ppmw] 449.7419621 258.501177 154.7963568 94.89542051 59.02479868 37.07167721 23.45258097 14.92938642
MeOH out [ppmw] 1829.993365 1084.814414 670.9935117 424.9605146 272.8785084 176.6766004 114.9452838 74.99016633
pH bottom tray 7.367803 7.278316 7.166382 7.058951 6.954706 6.852524 6.751812 6.652326
purification CO2 0.999543093 0.999989509 0.999999772 0.999999995 1 1 1 1
purification NH3 0.784893125 0.876361592 0.925962522 0.954612513 0.971769056 0.982269005 0.988782876 0.992859431
purification MeOH 0.775189277 0.866732898 0.9175699 0.947794521 0.966477466 0.978295662 0.985879221 0.99078762
No. trays 10 11 12 13 14 15 16 17
CO2 out [ppmw] 7.90807E-15 6.46079E-17 5.21593E-19 4.12742E-21 3.23368E-23 2.53213E-25 1.97161E-27 1.51632E-29
NH3 out [ppmw] 9.555236288 6.150035303 3.976786818 2.584654812 1.687018478 1.104229694 0.725052378 0.47821728
MeOH out [ppmw] 49.01284256 32.06310507 20.98943532 13.74357128 9.000840888 5.895892209 3.862221912 2.529890857
pH bottom tray 6.5539 6.45664 6.360471 6.265753 6.172743 6.081861 5.994309 5.91193
purification CO2 1 1 1 1 1 1 1 1
purification NH3 0.99542983 0.997058502 0.998097944 0.998763787 0.999193117 0.999471858 0.999653215 0.999771274
purification MeOH 0.993978878 0.996061117 0.997421493 0.998311632 0.998894266 0.999275702 0.999525534 0.999689208
No. trays 18 19 20 IN [ppmw]
CO2 out [ppmw] 1.16118E-31 8.92684E-34 6.78846E-34 CO2 out 9097.474148
NH3 out [ppmw] 0.316432484 0.209622387 0.139523918 NH3 out 2090.783765
MeOH out [ppmw] 1.657190865 1.085615729 0.711128912 MeOH out 8140.151598
pH bottom tray 5.836425 5.769461 5.71335
purification CO2 1 1 1
purification NH3 0.999848654 0.99989974 0.999933267
purification MeOH 0.999796418 0.999866634 0.999912639
d
5 trays
S:F 0.2 0.25 0.3 0.34 0.3824 IN [ppmw]
Total steam [kg/h] 6380 7975 9570 10846 12349 CO2 out 9097.474148
CO2 out [ppmw] 0.003864048 0.000859806 0.000238304 0.0001059 4.26932E-05 NH3 out 2090.783765
NH3 out [ppmw] 665.2173679 390.9451243 225.7598941 152.7770813 94.89542051 MeOH out 8140.151598
MeOH out [ppmw] 2792.280498 1688.962695 995.1287455 679.6717519 424.9605146
pH bottom tray 7.48848 7.370511 7.249404 7.163536 7.058951
purification CO2 0.999999575 0.999999905 0.999999974 0.999999988 0.999999995
purification NH3 0.681833493 0.813015037 0.892021405 0.92692832 0.954612513
purification MeOH 0.656974386 0.79251459 0.877750588 0.916503797 0.947794521
10 trays
S:F 0.2 0.25 0.3 0.34 0.3824
Total steam [kg/h] 6380 7975 9570 10846 12349
CO2 out [ppmw] 2.06257E-09 2.91416E-11 5.98026E-13 7.69696E-14 7.90807E-15
NH3 out [ppmw] 544.6058635 208.0588508 64.30452431 27.07343111 9.555236288
MeOH out [ppmw] 2412.991213 1015.582859 332.5251321 141.3262723 49.01284256
pH bottom tray 7.444117 7.231568 6.973569 6.78342 6.5539
purification CO2 1 1 1 1 1
purification NH3 0.739520713 0.900487628 0.969243819 0.987051061 0.99542983
purification MeOH 0.703569254 0.875237845 0.959150007 0.982638373 0.993978878
15 trays
S:F 0.2 0.25 0.3 0.34 0.3824
Total steam [kg/h] 6380 7975 9570 10846 12349
CO2 out [ppmw] 1.52949E-16 1.2366E-19 2.08369E-22 7.2184E-24 2.53213E-25
NH3 out [ppmw] 513.9106819 137.7388262 21.24332863 5.410078031 1.104229694
MeOH out [ppmw] 2340.300944 758.9026868 128.8322964 32.0394682 5.895892209
e
pH bottom tray 7.431301 7.140966 6.730064 6.428352 6.081861
purification CO2 1 1 1 1 1
purification NH3 0.754201898 0.93412096 0.989839538 0.997412416 0.999471858
purification MeOH 0.712499096 0.906770448 0.984173231 0.996064021 0.999275702
20 trays
S:F 0.2 0.25 0.3 0.34 0.3824
Total steam [kg/h] 6380 7975 9570 10846 12349
CO2 out [ppmw] 7.61696E-24 3.92569E-28 4.48623E-32 5.01467E-34 6.78846E-34
NH3 out [ppmw] 504.5540612 101.2145969 7.386518239 1.13930696 0.139523918
MeOH out [ppmw] 2325.666043 625.7290988 51.89185523 7.349102664 0.711128912
pH bottom tray 7.427254 7.073337 6.497098 6.088494 5.71335
purification CO2 1 1 1 1 1
purification NH3 0.758677071 0.951590117 0.996467106 0.999455081 0.999933267
purification MeOH 0.714296962 0.923130535 0.993625198 0.999097179 0.999912639
Stripper with reboiler
5 trays
Reboil Ratio 0.95 0.925 0.9 0.875 0.85 0.825 0.8 0.775 0.75
NH3 out [ppmw] 1.90E+03 1.69E+03 1.45E+03 1.22E+03 1.00E+03 8.06E+02 6.36E+02 4.96E+02 3.84E+02
CO2 out [ppmw] 1.32E+01 2.82E-01 6.65E-02 2.35E-02 9.50E-03 4.18E-03 1.96E-03 9.65E-04 5.01E-04
MeOH out [ppmw] 7.41E+03 6.64E+03 5.79E+03 4.95E+03 4.13E+03 3.38E+03 2.71E+03 2.14E+03 1.68E+03
pH bottom tray 7.687188 7.697648 7.664225 7.580635 7.531502 7.531502 7.423268 7.366466 7.366499
purification NH3 0.091761116
0.194019594
0.306243735
0.416449847
0.520809703
0.615027791
0.696147879
0.763173524
0.816751836
purification CO2 0.99852175 0.99996855 0.99999257 0.99999737 0.99999893 0.99999953 0.99999978 0.99999989 0.99999994
purification MeOH 0.08473267 0.17918827 0.28456808 0.38914918 0.48989539 0.58304948 0.66546039 0.73541843 0.79274075
Reboil Ratio 0.725 0.7 0.675 0.65 0.625 0.6 IN [ppmw]
NH3 out [ppmw] 2.96E+02 2.29E+02 1.78E+02 1.38E+02 1.09E+02 8.60E+01 NH3 out 2094.49189
f
CO2 out [ppmw] 2.72E-04 1.54E-04 9.14E-05 5.60E-05 3.56E-05 2.34E-05 CO2 out 8956.94456
MeOH out [ppmw] 1.31E+03 1.02E+03 7.95E+02 6.20E+02 4.89E+02 3.87E+02 MeOH out 8095.45151
pH bottom tray 7.309268 7.252451 7.196628 7.141449 7.088604 7.037442
purification NH3 0.85856935 0.89071691 0.91520885 0.93403597 0.94814256 0.95892143
purification CO2 0.99999997 0.99999998 0.99999999 0.99999999 0.99999999 0.99999999
purification MeOH 0.83844573 0.87422282 0.90185162 0.92341993 0.93964871 0.95213564
10 trays
Reboil Ratio 0.95 0.925 0.9 0.875 0.85 0.825 0.8 0.775 0.75
NH3 out [ppmw] 1.90E+03 1.69E+03 1.44E+03 1.18E+03 9.20E+02 6.65E+02 4.43E+02 2.72E+02 1.57E+02
CO2 out [ppmw] 1.93E+00 1.44E-05 1.14E-06 1.58E-07 2.47E-08 2.83E-09 3.86E-10 4.48E-11 6.60E-12
MeOH out [ppmw] 7.41E+03 6.64E+03 5.76E+03 4.84E+03 3.89E+03 2.94E+03 2.06E+03 1.33E+03 7.97E+02
pH bottom tray 7.750979 7.698572 7.617877 7.561203 7.488758 7.488758 7.398174 7.290351 7.043246
purification NH3 0.09105887 0.19511405 0.31260715 0.43514813 0.56077129 0.68231181 0.78870158 0.87029681 0.92494542
purification CO2 0.99978423 0.99999999 1 1 1 1 1 1 1
purification MeOH 0.08406159 0.17982251 0.28874538 0.40202912 0.51924621 0.63653092 0.74567968 0.83618647 0.90159790
Reboil Ratio 0.725 0.7 0.675 0.65 0.625 0.6
NH3 out [ppmw] 8.83E+01 4.94E+01 2.78E+01 1.61E+01 9.59E+00 5.86E+00
CO2 out [ppmw] 1.16E-12 1.77E-13 4.92E-14 1.53E-14 5.17E-15 1.79E-15
MeOH out [ppmw] 4.58E+02 2.59E+02 1.45E+02 8.31E+01 4.84E+01 2.88E+01 IN [ppmw]
pH bottom tray 6.915647 6.915647 6.789433 6.669247 6.554604 6.446028 NH3 out 2094.49189
purification NH3 0.95783671 0.97641438 0.98671587 0.99230341 0.99542332 0.99720151 CO2 out 8956.94456
purification CO2 1 1 1 1 1 1 MeOH out 8095.45151
purification MeOH 0.94342261 0.96804153 0.98207187 0.98973118 0.99401730 0.99644602
15 trays
Reboil Ratio 0.95 0.925 0.9 0.875 0.85 0.825 0.8 0.775 0.75
NH3 out [ppmw] 1.90E+03 1.69E+03 1.44E+03 1.18E+03 9.07E+02 6.28E+02 3.72E+02 1.83E+02 7.67E+01
CO2 out [ppmw] 9.53E-01 7.42E-10 2.40E-12 7.11E-14 3.23E-15 1.92E-16 7.08E-18 2.84E-19 8.03E-21
g
MeOH out [ppmw] 7.41E+03 6.64E+03 5.76E+03 4.84E+03 3.87E+03 2.86E+03 1.86E+03 1.01E+03 4.55E+02
pH bottom tray 7.757169 7.698578 7.662395 7.617277 7.55794 7.47595 7.360069 7.203051 7.012466
purification NH3 0.09107235 0.19512266 0.31283503 0.43668202 0.56718057 0.70017224 0.82222452 0.9127967 0.96336278
purification CO2 0.999893549
1 1 1 1 1 1 1 1
purification MeOH 0.08406825 0.17982211 0.28883831 0.40270352 0.52243319 0.64713643 0.77056496 0.87561955 0.94384249
Reboil Ratio 0.725 0.7 0.675 0.65 0.625 0.6 IN [ppmw]
NH3 out [ppmw] 3.02E+01 1.19E+01 5.00E+00 2.25E+00 1.09E+00 5.67E-01 NH3 out 2094.49189
CO2 out [ppmw] 4.35E-22 3.39E-23 3.11E-24 5.39E-25 1.13E-25 2.63E-26 CO2 out 8956.94456
MeOH out [ppmw] 1.84E+02 7.11E+01 2.83E+01 1.16E+01 4.93E+00 2.18E+00 MeOH out 8095.45151
pH bottom tray 6.807666 6.60266 6.411132 6.078531 5.944932 5.944932
purification NH3 0.98557026 0.99430995 0.99761089 0.99892730 0.99948101 0.99972938
purification CO2 1 1 1 1 1 1
purification MeOH 0.97727877 0.99121570 0.99650221 0.99856612 0.99939074 0.99973130
20 trays
Reboil Ratio 0.95 0.925 0.9 0.875 0.85 0.825 0.8 0.775 0.75
NH3 out [ppmw] 1.90E+03 1.69E+03 1.44E+03 1.18E+03 9.04E+02 6.16E+02 3.39E+02 1.35E+02 4.02E+01
CO2 out [ppmw] 2.20E-01 3.73E-14 1.35E-18 1.70E-20 3.98E-22 1.09E-23 1.57E-25 1.42E-27 6.51E-30
MeOH out [ppmw] 7.42E+03 6.64E+03 5.76E+03 4.84E+03 3.86E+03 2.84E+03 1.78E+03 8.41E+02 2.79E+02
pH bottom tray 7.761832 7.69857 7.662392 7.617223 7.557356 7.471756 7.339749 7.137162 6.87044
purification NH3 0.09105614 0.19515154 0.31284357 0.43681885 0.56832064 0.70581679 0.83790821 0.93540248 0.98080766
purification CO2 0.99997542 1 1 1 1 1 1 1
purification MeOH 0.08405018 0.17984803 0.28884018 0.40273836 0.52276396 0.64931747 0.77971179 0.89616564 0.96549223
Reboil Ratio 0.725 0.7 0.675 0.65 0.625 0.6 IN [ppmw]
NH3 out [ppmw] 1.10E+01 3.19E+00 1.09E+00 4.33E-01 1.99E-01 1.01E-01 NH3 out 2094.49189
CO2 out [ppmw] 9.00E-32 2.79E-33 1.29E-33 1.39E-33 1.96E-33 3.25E-33 CO2 out 8956.94456
MeOH out [ppmw] 7.66E+01 2.01E+01 5.57E+00 1.63E+00 5.03E-01 1.65E-01 MeOH out 8095.45151
h
pH bottom tray 6.584478 6.312093 6.078344 5.893022 5.761544 5.676517
purification NH3 0.99476046 0.99847596 0.99948147 0.99979337 0.99990509 0.99995176
purification CO2 1 1 1 1 1 1
purification MeOH 0.99053314 0.99752313 0.99931240 0.99979906 0.99993785 0.99997966
Reboil ratio = 0.7
Number of trays 5 6 7 8 9 10 11 12 13
NH3 out [ppmw] 2.08E+02 1.50E+02 1.10E+02 8.08E+01 5.99E+01 4.46E+01 3.34E+01 2.51E+01 1.89E+01
CO2 out [ppmw] 1.35E-04 3.50E-06 8.55E-08 1.05E-09 1.23E-11 1.40E-13 1.55E-15 1.67E-17 1.74E-19
MeOH out [ppmw] 1.01E+03 7.56E+02 5.71E+02 4.35E+02 3.33E+02 2.56E+02 1.97E+02 1.52E+02 1.18E+02
pH bottom tray 7.23103 7.159659 7.090749 7.023696 6.958009 6.893378 6.829605 6.766574 6.704232
purification NH3 0.89108115 0.92127121 0.94246919 0.95760772 0.96856671 0.97657775 0.98247573 0.98684127 0.99008589
purification CO2 0.99999998 1 1 1 1 1 1 1 1
purification MeOH 0.87493560 0.90662984 0.92951863 0.94635857 0.95893798 0.96843787 0.97566882 0.98120414 0.98545882
Number of trays 14 15 16 17 18 19 20 IN [ppmw]
NH3 out [ppmw] 1.43E+01 1.08E+01 8.25E+00 6.30E+00 4.84E+00 3.73E+00 2.89E+00 NH3 out 1905.58179
CO2 out [ppmw] 1.78E-21 1.78E-23 1.75E-25 1.69E-27 1.61E-29 1.50E-31 1.39E-33 CO2 out 8477.56234
MeOH out [ppmw] 9.12E+01 7.07E+01 5.48E+01 4.25E+01 3.30E+01 2.56E+01 1.99E+01 MeOH out 8100.56318
pH bottom tray 6.642569 6.581615 6.521432 6.462108 6.403757 6.346517 6.290549
purification NH3 0.99250553 0.99431510 0.99567194 0.99669186 0.99746048 0.99804125 0.99848135
purification CO2 1 1 1 1 1 1 1
purification MeOH 0.98873881 0.99127270 0.99323311 0.99475141 0.99592816 0.99684063 0.99754842
20 trays
Reboil Ratio 0.8 0.79 0.78 0.77 0.76 0.75 0.74 0.73 0.72
NH3 out [ppmw] 3.05E+02 2.19E+02 1.49E+02 9.56E+01 5.90E+01 3.58E+01 2.13E+01 1.26E+01 7.58E+00
CO2 out [ppmw] 6.65E-26 1.38E-26 2.18E-27 2.92E-28 3.36E-29 3.88E-30 6.68E-31 1.35E-31 2.85E-32
MeOH out [ppmw] 1.76E+03 1.35E+03 9.75E+02 6.67E+02 4.32E+02 2.71E+02 1.63E+02 9.59E+01 5.66E+01
i
pH bottom tray 3.77985465 4.10132626 4.2633571 4.42564822 4.42564822 4.58867604 4.75127494 4.91372309 5.07661621
purification NH3 0.84012172 0.88482586 0.92206876 0.94981672 0.96902028 0.98122366 0.98882921 0.99339086 0.99601983
purification CO2 1 1 1 1 1 1 1 1 1
purification MeOH 0.78329328 0.83323591 0.87960763 0.91762803 0.94666359 0.96659219 0.97985520 0.98815685 0.99301372
Reboil Ratio 0.71 0.7 0.69 0.68 0.67 IN [ppmw]
NH3 out [ppmw] 4.64E+00 2.88E+00 1.85E+00 1.21E+00 8.08E-01 NH3 out 1905.58179
CO2 out [ppmw] 6.39E-33 1.83E-33 4.45E-33 1.61E-33 6.18E-34 CO2 out 8477.56234
MeOH out [ppmw] 3.34E+01 1.95E+01 1.17E+01 7.04E+00 4.26E+00 MeOH out 8100.56318
pH bottom tray 5.23953745 5.40216423 5.56524963 5.72836188 5.89130515
purification NH3 0.99756529 0.99848649 0.99903011 0.99936686 0.99957578
purification CO2 1 1 1 1 1
purification MeOH 0.99588128 0.99759202 0.99855901 0.99913069 0.99947448
Steam stripper with pH adjustment
pH 9 on tray 1 column 2
S:F 1 0.203534996 0.17222192 0.140908843 0.109595767 In stripper 1 Valid for all runs
S:F 2 0.199120988 0.229566374 0.259953068 0.290306808 CO2 [ppmw] 9097.472596
Total steam [kg/h] 13000 13000 13000 13000 NH3 [ppmw] 2090.783497
S:F 1/S:F 2 1.022167466 0.750205341 0.542054935 0.37751704 MeOH [ppmw] 8140.150441
Stripper 1
CO2 out [ppmw] 0.003705335 1.13E-02 3.94E-02 0.200493827
NH3 out [ppmw] 656.517521 888.1117982 1163.193031 1467.922626
MeOH out [ppmw] 2758.045024 3658.05752 4698.666991 5831.499817
Stripper 2
CO2 out [ppmw] 6.96185E-05 7.07E-05 8.12E-05 1.44E-04
NH3 out [ppmw] 122.8894975 94.25914548 63.63540278 39.25164976
MeOH out [ppmw] 607.517736 488.7836543 345.0622826 220.3301786
purification CO2 0.999999992 0.999999992 0.999999991 0.999999984
purification NH3 0.941223232 0.954916831 0.969563849 0.981226344
j
purification MeOH 0.925367751 0.939953978 0.95760984 0.972932911
pH 10 on tray 1 column 2
S:F 1 0.203534996 0.17222192 0.140908843 0.109595767
S:F 2 0.199133372 0.229566093 0.25995275 0.290306365
Total steam [kg/h] 13000 13000 13000 13000
S:F 1/S:F 2 1.022103901 0.750206259 0.542055598 0.377517617
Stripper 1
CO2 out [ppmw] 0.003825536 1.13E-02 3.94E-02 0.200502077
NH3 out [ppmw] 656.2409624 888.1195874 1163.200171 1467.930773
MeOH out [ppmw] 2758.195104 3658.080589 4698.694884 5831.530633
Stripper 2
CO2 out [ppmw] 0.003632986 1.06E-02 3.68E-02 0.185781598
NH3 out [ppmw] 93.73211989 65.84945588 41.26596018 23.92886774
MeOH out [ppmw] 538.8230973 416.5434975 283.5694778 175.1650723
purification CO2 0.999999601 0.999998833 0.999995957 0.999979579
purification NH3 0.955168902 0.96850489 0.98026292 0.988555071
purification MeOH 0.93380674 0.948828526 0.965164099 0.978481347
pH 10.5 on tray 1 column 2
S:F 1 0.203534996 0.17222192 0.140908843 0.109595767 0.121338171 0.080944302 0.121338171 0.080944302
S:F 2 0.199138618 0.229579304 0.259953068 0.290306808 0.248370309 0.288661298 0.217797527 0.258001575
Total steam [kg/h] 13000 13000 13000 13000 12000 12000 11000 11000
S:F 1/S:F 2 1.022076971 0.750163089 0.542054935 0.37751704 0.488537343 0.280412728 0.557114546 0.31373569
Stripper 1
CO2 out [ppmw] 0.003513081 0.010967616 0.039393594 0.200494133 0.098633734 4.373238599 0.100278261 4.373189543
NH3 out [ppmw] 655.5762796 887.6846009 1163.193949 1467.922626 1350.598305 1721.627798 1348.502638 1721.627492
MeOH out [ppmw] 2752.000424 3655.563701 4698.67005 5831.499817 5398.12767 6767.401233 5391.761965 6767.398167
Stripper 2
CO2 out [ppmw] 0.003489761 0.010901025 0.039217083 0.199789419 0.098239762 4.343710536 0.099853871 4.343800506
k
NH3 out [ppmw] 3.963884608 1.78516178 0.8317811 0.412414088 1.547245723 0.592450292 4.918255824 1.693565631
MeOH out [ppmw] 169.0208613 98.35965656 53.21596345 28.50963451 88.31185362 37.90116884 216.2275783 92.56435983
purification CO2 0.999999616 0.999998802 0.999995689 0.999978039 0.999989201 0.999522537 0.999989024 0.999522527
purification NH3 0.998104115 0.999146176 0.999602168 0.999802747 0.999259968 0.999716637 0.997647649 0.999189985
purification MeOH 0.97923615 0.987916727 0.993462533 0.996497653 0.989151078 0.995343923 0.973436906 0.988628667
S:F 1 0.121338171 0.080944302
S:F 2 0.187239699 0.227341781
Total steam [kg/h] 10000 10000
S:F 1/S:F 2 0.648036562 0.356046749
Stripper 1
CO2 out [ppmw] 0.100200293 4.373568383
NH3 out [ppmw] 1351.011614 1721.62727
MeOH out [ppmw] 5399.124727 6767.399158
Stripper 2
CO2 out [ppmw] 0.099736908 4.344189891
NH3 out [ppmw] 16.88808174 5.309882201
MeOH out [ppmw] 496.3208229 228.3257906
purification CO2 0.999989037 0.999522484
purification NH3 0.991922606 0.997460339
purification MeOH 0.939028053 0.971950667
pH 11 on tray 1 column 2
S:F 1 0.203534996 0.17222192 0.140908843 0.109595767 0.121338171 0.080944302 0.121338171 0.080944302
S:F 2 0.199141486 0.229566374 0.259953068 0.290306808 0.248370309 0.288661298 0.217809445 0.258001496
Total steam [kg/h] 13000 13000 13000 13000 12000 12000 11000 11000
S:F 1/S:F 2 1.022062254 0.750205341 0.542054935 0.37751704 0.488537343 0.280412728 0.557084062 0.313735787
Stripper 1
CO2 out [ppmw] 0.003826887 0.011275137 0.039393288 0.200493827 0.098633734 4.373238599 0.10019479 4.373562251
NH3 out [ppmw] 655.9990147 888.1142469 1163.193031 1467.922626 1350.598305 1721.627798 1351.009474 1721.62727
l
MeOH out [ppmw] 2757.364482 3658.05752 4698.666991 5831.499817 5398.12767 6767.401233 5399.118613 6767.399158
Stripper 2
CO2 out [ppmw] 0.003353586 0.00989855 0.034622701 0.176407448 0.086787928 3.839185762 0.088112733 3.842097863
NH3 out [ppmw] 0.000125847 4.50018E-05 1.84414E-05 8.24862E-06 3.37641E-05 1.10965E-05 0.00012014 3.27946E-05
MeOH out [ppmw] 3.941622478 1.578807348 0.685951766 0.317337228 1.18573927 0.407151092 3.721944896 1.114884393
purification CO2 0.999999631 0.999998912 0.999996194 0.999980609 0.99999046 0.999577994 0.999990315 0.999577674
purification NH3 0.99999994 0.999999978 0.999999991 0.999999996 0.999999984 0.999999995 0.999999943 0.999999984
purification MeOH 0.99951578 0.999806047 0.999915732 0.999961016 0.999854334 0.999949982 0.999542767 0.999863039
S:F 1 0.121338171 0.080944302
S:F 2 0.187239699 0.227341781
Total steam [kg/h] 10000 10000
S:F 1/S:F 2 0.648036562 0.356046749
Stripper 1
CO2 out [ppmw] 0.10019479 4.373562251
NH3 out [ppmw] 1351.009474 1721.62727
MeOH out [ppmw] 5399.118613 6767.399158
Stripper 2
CO2 out [ppmw] 0.088207123 3.844974254
NH3 out [ppmw] 0.000502485 0.00011065
MeOH out [ppmw] 12.98826996 3.378574603
purification CO2 0.999990304 0.999577358
purification NH3 0.99999976 0.999999947
purification MeOH 0.998404419 0.999584949
m
C. Economy calculations The economical calculations are based on the Ulrich method. [78]
Material for the ion exchangers
Purchase equipment cost for a process vessel, CP, with the current dimensions was evaluated to
8 000 USD. From tables and diagrams [78] the pressure and material factor and the installation
factor, FaBM, could be read to 12.4 and 23.8.
With all the factors the price of a vessel, CBM, was estimated to 190 400 USD. [78]
𝐶𝐵𝑀 = 𝐶𝑃 ∙ 𝐹𝐵𝑀𝑎 (C-eq1)
The update factor is 1.809 estimated from net price index: [79]
𝑓 = 𝐼𝑁𝑃 2009
𝐼𝑁𝑃 1984 (C-eq2)
The updated price for a vessel is 344 500 USD and the total cost of five vessels are estimated to
1 722 500 USD, not including installation cost.
The price of the packing material is 133 USD/MT [77] and the total cost is 1 800 USD.
𝑇𝑜𝑡𝑎𝑙 𝑐𝑜𝑠𝑡𝑝𝑎𝑐𝑘𝑖𝑛𝑔 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙 = 𝑚𝑐𝑙𝑖𝑛𝑜𝑝𝑡𝑖𝑙𝑜𝑙𝑖𝑡𝑒 ∙ 𝑃𝑟𝑖𝑐𝑒 𝑜𝑓 𝑝𝑎𝑐𝑘𝑖𝑛𝑔 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙 (C-eq3)
This means that the total cost of five ion-exchangers including the packing material is 1 724 300 USD,
not including installing costs.
𝑇𝑜𝑡𝑎𝑙 𝑐𝑜𝑠𝑡𝑖𝑜𝑛−𝑒𝑥𝑐𝑎𝑛𝑔𝑒𝑟 = 𝑃𝑟𝑖𝑐𝑒 𝑜𝑓 𝑝𝑎𝑐𝑘𝑖𝑛𝑔 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙 + 𝑃𝑟𝑖𝑐𝑒 𝑜𝑓 𝑝𝑟𝑜𝑐𝑒𝑠𝑠 𝑣𝑒𝑠𝑠𝑒𝑙𝑠 (C-eq4)
Material for the strippers
Purchase cost for each tray, CP, with the current dimensions was evaluated to 280 USD from a
diagram. [79] The diameter of the trays is given from the simulation in Aspen as one meter. The
material factor for stainless steel is FBM=2.0 and the tray efficiency coefficient is estimated to η = 0.6.
With all these factors the prices of the trays were estimated with the equation below:
𝐶𝐵𝑀 = 𝐶𝑝 ∙ 𝐹𝐵𝑀 ∙ 𝑁𝐴𝑐𝑡 ∙ 𝑓𝑞 (C-eq5)
Where NAct is the number of trays and fq is a price factor that decreases with the number of trays.
The results are shown in table C-1.
Table C-1 Tray cost
Ideal trays Real trays (NAct) fq CBM [USD] Price [USD]
5 9 1.7 8568 15 500
10 17 1.2 11424 20 700
15 25 1.0 14000 25 300
20 34 1.0 19040 34 500
n
The height of the distillation tower, H, is calculated using the distance between two trays, h, the
number of trays, n, and 1.5 meters extra space above the highest and below the lowest tray, as seen
in equation below.
𝐻 = ∙ 𝑛 − 1 + 3 (C-eq6)
Besides this dimensions some material parameters are needed to calculate the cost of the vessel. The
chosen material is stainless steel which gives a material factor 𝐹𝑀 of 4.0. The pressure in the column
gives a pressure factor 𝐹𝑝 of 2.8. These together give the Bare module factor 𝐹𝐵𝑀𝛼 as 21.5. [78]
The purchased equipment cost 𝐶𝑝 is calculated from the dimensions given and allows 𝐶𝐵𝑀 to be
calculated as.
𝐶𝐵𝑀 = 𝐶𝑝 ∙ 𝐹𝐵𝑀𝛼 (C-eq7)
The result is summarized in table C-2, where the last column is including the cost for the trays. Note
that this cost is without the price of installation.
Table C-2 Vessel cost
Ideal trays Real trays Height CP [USD] CBM [USD] Price [USD] Total material
cost [USD]
5 9 7.9 15 000 323 000 584 000 600 000
10 17 12.8 22 000 473 000 856 000 877 000
15 25 17.6 31 000 667 000 1 207 000 1 232 000
20 34 23.1 41 000 882 000 1 596 000 1 631 000
A sensitivity analysis of the price dependency of the tray efficiency shows that up to 10-15 ideal trays
the price is quite constant. It also shows that if the tray efficiency is decreased below 0.5-0.6, price is
increased more distinctly.
Figure C-1 Sensitivity analysis of the price dependent on tray efficiency, based on Ulrich method [78]
0
500 000
1 000 000
1 500 000
2 000 000
2 500 000
0.40 0.50 0.60 0.70 0.80
Pri
ce [
USD
]
Tray efficiency
5 ideal trays
10 ideal trays
15 ideal trays
20 ideal trays
o
Total cost
To calculate the actual cost for the process supplement charge factor has to be added to compensate
for both contract by tender work for installation, unexpected costs and cost for help equipment. This
is done by the equation below.
𝐾 = 𝐶𝐵𝑀 𝑖𝑛𝐼=1 ∙ 𝑓𝑐𝑜𝑛/𝑢𝑛 ∙ 𝑓𝑒𝑙𝑝 (C-eq8)
As a rule of thumb according to Ulrich´s method the factor for contract by tender and unexpected
work is about 𝑓𝑐𝑜𝑛 /𝑢𝑛 ≈ 1.15. The factor for help equipment could be up to 1.25, but because the
plant is already existing and this is only a new part of it the factor is decreased to 𝑓𝑒𝑙𝑝 ≈ 1.15.
If the costs that are calculated in the section above is correct with these supplement charge factors
the total costs for the equipment becomes as seen in table C-3.
Table C-3 Total cost for each equipment
Equipment Total cost K [USD]
Stripper with 9 trays 793 500
Stripper with 17 trays 1 160 000
Stripper with 25 trays 1 596 000
Stripper with 34 trays 2 111 000
5 ion exchangers 2 280 000
The price for caustic soda varies between 800 and 820 USD/MT for northwestern Europe, the report
is from the 7th of November 2008. [80] Using the mean value of the price gives the cost for different
steam to feed ratios as seen in table C-4.
p
Table C-4 Cost for caustic for different steam to feed ratios in stripper 1.
S:F ratio stripper 1 Amount of caustic [kg/h] Cost for caustic [USD/h]
0.080 132 107
0.11 113 91
0.12 104 84
0.14 89 72
0.17 68 55
0.20 50 41
q
D. Calculation method in Aspen To simulate the rest of the process Aspen Tech’s flow sheeting software Aspen Plus 2004.1 was used.
It is a powerful tool with extensive databases of pure components and phase equilibrium data for
conventional chemicals and electrolytes. The program needs a lot of settings to be made before
proper simulations can be achieved. The following will explain what settings was used in the
simulations and how they were set.
The template “Electrolytes with Metric Units” was used which makes it possible to handle ions in the
system. The template automatically specifies water as component and thereby also provides solvent
for the ions. To model the stripper a “RadFrac” unit was used which is a rigorous model for
simulating all types of multistage vapour-liquid fractionation operations.
When specifying the components used in the system all species but water is specified, that is: 𝐶𝐻4,
𝐶𝑂, 𝐶𝑂2, 𝐻2, 𝑁2, 𝑁𝐻3 and 𝐶𝐻3𝑂𝐻. To generate the ions the electrolyte wizard is used. In the wizard
all ionic components suggested except 𝑁𝐻2𝐶𝑂𝑂− is chosen, that is: 𝑁𝐻4+, 𝐻𝐶𝑂3
−, 𝐶𝑂32−, 𝐻3𝑂
+ and
𝑂𝐻−. To minimize unwanted interference 𝑁𝐻2𝐶𝑂𝑂− was deselected, this was supported by simple
simulations showing almost no trace of the component (3 ∙ 10−24 ppm) in the processed
condensate. The salt formation is deselected to minimize interference and this is sanctioned by that
the possible salts formed are readily soluble. The wizard also generates Henry Law components for
the following species: 𝐶𝑂2, 𝑁𝐻3, 𝐻2 and 𝑁2, which thereby uses Henry’s Law to calculate the
solubility of these components.
The property method used is ELECNRTL which is suggested by the electrolyte wizard. That is an
electrolyte NRTL model (Non-Random Two Liquid activity coefficient model) with Redlich-Kwong
equation of state valid for aqueous and mixed solvent applications. Another property method
possible for this simulation is the Pitzer method. [73] Since the databank for Pitzer parameters did
not include some of the ion pairs the Pitzer property method was rejected. [74]