June 2
PROJECT REPORT
2014
Design and simulation of CO2absorption and stripping section for removal of co2.
ACKNOWLEDGEMENT
All praises to Almighty ALLAH who gave us light in darkness and gave us understanding and ability to complete our report and all respects are for his Prophet MUHAMMAD (PBUH, on whom be ALLAH,S blessings and salutations) I would like to thank PROF. DR. MAHMOOD SALEEM for granting me the chance to pursue this Assignment in an environment that facilitated my learning. I take immense pleasure in thanking our worthy teacher for valuable help regarding this assignment. Abid Hussain Roll No. PG-M10-20 ICET University of the Punjab new campus Lahore.
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PLANT DESIGN ASSIGNMENT REPORTE ON DESIGNING OF SYNTHESIS GAS PURIFICATION (CO2 removal) SECTION USING ACTIVATED-MDEA
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PLANT DESIGN MID TERM ASSIGNMENT REPORTE ON
DESIGNING OF SYNTHESIS GAS PURIFICATION (CO2 removal) SECTION USING ACTIVATED-MDEA
BSc. Engineering (7th Semester)
Submitted By:
Abid Hussain
Roll No.
PG-M10-20
Supervised by:
Prof. Dr. Mahmood Saleem
Institute of Chemical Engineering and Technology Faculty of Engineering & Technology University of the Punjab New Campus Lahore
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ABSTRACT: Ammonia is an important and useful for the production of urea. Since then it has been widely used in the production of other chemicals, products and fertilizers specially. The synthesis gas stream leaving the low temperature shift converter contains approximately 18.4 mole% Carbon dioxide on a dry gas basis. It is essential to remove all CO2 from the synthesis gas before entering the ammonia synthesis loop. Here we will use activated MDAE as an absorbent for CO2 removal. Solution in our process. Hence in the first chapter of this report I have given the introduction and contextual of CO2 gas. Removal process using aMDEA as solvent. The second chapter includes the detailed process description of purification section, equipment working , feed and product composition and conditions. The 3rd chapter contained detail material and energy balance calculation using chemcad for all the process equipment to design Syn. Gas purification section of ammonia plant. I had 6 week training at FAUJI FERTILIZER COMPANY ,SADIQ ABAD, RAHIM YAR KHAN . So the design data I used in my calculation is courtesy of FAUJI FERTILIZER COMPANY. I design the process in following steps.1st , draw its flow sheet , list down all required equipment and then I performed material and energy balance of the equipment and of the whole process using CHEMCAD. I hope that the reader will find the information contained to be useful.
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Contents Chapter 1................................................................................................................................................. 9 1.1 INTRODUCTION ................................................................................................................................. 9 1.2Synthesis Gas purification process:- .................................................................................................. 9 1.3Background of purification processes:- ............................................................................................. 9 1.4 ......................................................................................................................................................... 10 CHAPTER 2 ............................................................................................................................................ 11 PROCESS DESCRIPTION ..................................................................................................................... 11 2.1Absorption Column:- ................................................................................................................ 11 2.2Absorbent Regeneration (CO2stripping Section:- ................................................................... 12 2.2) Low pressure flashing column:- ............................................................................................. 12 2.3) Stripper :- .............................................................................................................................. 13 CHAPTER 3 ............................................................................................................................................ 14 MATERIAL AND ENERGY BALANCE ................................................................................................... 14 CHAPTER NO.4 ...................................................................................................................................... 17 4.1 Absorber Design: ......................................................................................................................... 17 Select between Plate & Packed column: ...................................................................................... 17 Factors affecting the absorption Column : ................................................................................... 18 Foaming......................................................................................................................................... 18 Entrainment .................................................................................................................................. 18 Weeping/Dumping:....................................................................................................................... 19 Flooding......................................................................................................................................... 19 State of trays & Packing: ............................................................................................................... 19 Column Diameter: ......................................................................................................................... 19 Standard Design steps:.................................................................................................................. 19 1.Calculation of theoretical number of stages:35 .............................................................................. 21 Calculation of Diameter of Column:37 ............................................................................................... 22 Calculation of Pressure drop: ............................................................................................................ 25 Down comer Design: ......................................................................................................................... 26 Entrainment Calculation: .................................................................................................................. 27 Calculation of Height of Column: ...................................................................................................... 28 4.2 Stripper Design:............................................................................................................................... 28 6
Stripper ............................................................................................................................................. 28 Stripping Phenomenon: ................................................................................................................ 28 Stripping Agents: ........................................................................................................................... 28 Types of Stripper: .............................................................................................................................. 28 ............................................................................................................. Error! Bookmark not defined. Standard Design Steps: ................................................................................................................. 29 Calculation of Weeping Point: ...................................................................................................... 32 Calculation of Pressure drop: ........................................................................................................ 34 Calculation of Height of Column: .................................................................................................. 36 4.5 MDEA Surge Drum: ......................................................................................................................... 37 Design Specifications .................................................................................................................... 37 CHAPTER NO.5 ...................................................................................................................................... 38 Dynamic Simmulations. .................................................................................................................... 38 CHAPTER NO.6 ...................................................................................................................................... 39 6.1 Plant Cost Estimation: ............................................................................................................... 39 6.2 Capital Investment: ................................................................................................................... 40 6.2.1 6.2.2
Direct costs:................................................................................................................... 40 Indirect costs: .................................................................................................................... 40
6.3
Types of Cost Estimation:...................................................................................................... 40
6.4
Methods of Estimating Capital Investment: ......................................................................... 41
6.5 Percentage of Delivered Equipment Cost: ................................................................................ 41 Cost of Absorber: .......................................................................................................................... 41 7.6 Direct Cost: 50 .................................................................................................................................. 44 7.7 In-Direct cost: .................................................................................................................................. 44 OPERATIONAL PROBLEMS. ................................................................................................................... 45 7.1
Problems occurring during operation: .................................................................................. 45
7.2
Foaming:................................................................................................................................ 45
Causes of Foaming: ....................................................................................................................... 45 Prevention of Foaming:................................................................................................................. 46 7.3 Corrosion: .................................................................................................................................. 46 Mechanism of Corrosion: .............................................................................................................. 47 Metods of Minimizing Corrosive Attacks: ..................................................................................... 47
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7.4
Chemical Losses54 .................................................................................................................. 47
7.5
Losses due to Volatility: ........................................................................................................ 48
7.6
Entrainment: ......................................................................................................................... 48
References: ........................................................................................................................................... 48
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Chapter 1 1.1 INTRODUCTION Absorption of CO2 from process gas because carbon dioxide gas present in synthesis gas is poison for ammonia synthesis catalyst. In this section synthesis gas is proceed to remove CO2 and CO, producing a high purity H2and N2. Bulk removal of carbon dioxide is talented by the use of an improved benefield low heat process which uses the four stage flash of the semi lean solution. to minimize external heat requirements. Final removal of remaining CO2 and CO. is accomplished by catalytically converting the CO2 to methane and water in the mathenator using hydrogen. The MDAE “low heat” process circulates an aqueous sol. containing a nominal 37% MDAE and 3 % piprazine . This solution chemically combines with CO 2 on the process gas but not significantly with the other voters. Additives are injected into the solvent to enhance the CO2absorption rate, inhibit corrosion and to control foaming.
1.2Synthesis Gas purification process:There is long history of different processes used for gas cleansing, here is given their name just , while MDAE method will be discussed in detail. 1) MEA (mono ethanol amine ) process 2) Benfield process (also called Hot process) 3) Activated MDEA process
1.3Background of purification processes:In the early days of ammonia manufacture, monoethanolamine (MEA) was frequently used for CO2 removal from the synthesis gas. Somewhat later, hot potassium carbonate (the Benfield, or Hot Pot process) was used, often in a split flow configuration described as a two‐ stage Benfield Low Heat process for energy conservation. In the last 20 years, a very substantial fraction of these plants have been retrofitted using BASF’s a-MDEA process.
Activated MDEA as solvent:N‐Methyldiethanolamine (MDEA) is a tertiary amine whose amino group is incapable of reacting with CO2. However, it is alkaline and so is an excellent sink for protons produce by CO2hydrolysis. Because it is non‐reactive, aqueous MDEA by itself absorbs CO2 far too slowly to be an effective solvent for giving ammonia synthesis gas. But when spiked with a relatively small attentiveness of piperazine, a diamine that reacts extraordinarily fast with
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CO2, the resulting blend is an excellent solvent for treating syngas and removing CO2 in the production of LNG. In this paper, we first present the results of a measurable study of the piperazine promotion of MDEA, specificsly the effects of piperazine to MDEA ratio, total amine strength, and the treating temperature on presentation of a typical ammonia syngas CO2 removal system. In a recent patent, Waner et al. (2009) proposed using the alkali metal salts of a number of tertiary amino acids, appropriately promoted with reactive amines such as MEA. Some of the potentially more intereting results of our study include the utility of operating at higher temperatures with lower rather than higher total amine concentrations, and the existence of operating boundaries that can lead to unstable operation when approached too closely. Following a discussion of amino acids and their mode of operation, we critically analyze the possibility of using the potasium salt of the tertiary amino‐acid dimethylglycine, promoted with piperazine as a syngas treating solvent. The results show that it is possible to treat syngas quite effectively using such a solvent but with much lower concentration of the piperazine promoter. Furtermore, results suggest that the cross‐exchanger commonly used as a heat integration tool in treating plants can be completely eliminated.
1.4 aMDEA Composition:MDEA=37% Piprazine=3% and remaining is water
Advantages of using a-MDEA:There are many advantages of using α-MDEA solution some of which most important are
More CO2 absorption.
No time for regeneration.
Low energy requirement for regeneration.
In CO2 removal section Amerel is used as antifoaming agent for a-MDEA solution.
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CHAPTER 2 PROCESS DESCRIPTION After the synthsis gas has been prepared it is purified off carbon dioxide and carbon monoxide to yield a high purity nitrogen hydrogen synthesis gas. The CO2 removal system consists of an Absober and a CO2 Stripper with the carbonate solution circulating in a closed loop between the two. 2.1Absorption Column:-
Process gas leaving the top of CO2 Absorber Feed Gas Separator is now called as raw synthesis gas which enters the CO2 absorber at the bottom. The internal packing at the bottom part, BED3 and BED2 , is replaced from IMTP#50 to structured packing. The lean αMDEA solution at 500C from the lean rich exchanger and lean cooler enters the absorber from the middle inlet. While the two stream flow conter-current through the absorber, the lean solution gradually absorbs the CO2 present in the feed gas ,and leaves the absorber bottom as rich solution stream at 480C to LP flash column. The treated gas at 500C exits the absorber from the top through CO2 absorber top knokout drum. Mist carry-over to the downstream equipment is minimized by packing in the previous lean absorption section, which is now performing as “demister”.
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Process Flow Sheet:Pure syn. gass
Sour gas
Flashing column E-9
stripper
P-9 P-12
E-10 E-12 E-13 E-11 E-14 E-15
E-17 E-3
reboilerP-8
P-15
Absorption column E-20
Rich MDAE
P-6
P-14
heater
P-10 P-4 P-7
E-18 E-19
Rich solution pump
Regenerated MDAE
Plate and frame heat exchanger
FIG: 2.1 Purification section flowsheet. 2.2Absorbent Regeneration (CO2stripping Section:-
Stripping of CO2 from a-MDEA solution, to regenerate solvent for recycling is carried out in this section. Regeneration takes place in following two steps.
Flashing in Low pressure flashing column and Stripping in (Stripper) 2.2) Low pressure flashing column:The rich soltion at 850C is flushed in the newly installed LP-Flash column at 0.9kg/cm2.Liquid from LP flash column is pumped through to lean-rich exchanger where it is heated before going to the stripper. In stipper after separation of CO2 from the α-MDEA solution, CO2 product vapor is recycled to the bottom part of LP flash column. CO2 product vapor from LP flash and CO2 vapors from stripper is cooled to 380C by direct contact cooling with quench or reflux water in a packed bed above the striping secton of the LP flash column, quench water is circulated by the flash column Quinch Pump 12
to the LP flash quench cooler . In this exchanger the quench water is heat is rejected to the cooling water, water condnsed from the CO2 product vapour during cooling is removed from the cooling circuit to fulfil the water make up requirements. After being cooled the 99% CO2 product passes through the demister pad, exists the column and is exported for use in urea plant.
The main advantages of using LP Flash column are:
Chemically unbound CO2 molecules removal.
No heat requirement like stripper.
2.3)
Stripper :-
In LP flash column only chemical unbound molecules of CO2 are removed, chemically bound molecules are removed by using CO2 stripper. The a-MDEA solution containing chemically bound CO2 molecules exiting from bottom of LP flash column is sent to middle section of
stripper being pumped by rich solution
pump.Steam for the purpose of striping is produced in steam generator and provision is made to control steam supply from SH and SL heaters Regenerated solution leaves and goes to Solution tank by gravity flow is used as holding tank to provide residence time . The lean solution at 1190C from 117-F is then cooled down sequentially in lean-rich exchanger and lean cooler to 50% with concentration, prior to recycle into the absorber. Acid off gas and 114-C leaves striper from the top at 1040C is directed to the bottom part of LP flash column. CO2 gas separated from striper at 50ˆC is pass through heat exchanger and cooled to 38C to remove any entained water. Hence we obtained 99% CO2 which is transported to urea section. While absorber overhead gas, purified Syn. Gas , containing approx. 1000ppm of CO2 is disengaged of any entrained liquid and preheated to about 316C , in methanator effluent exchanger and finally sent to Methanator for further purification.
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CHAPTER 3 MATERIAL AND ENERGY BALANCE
FlowDiagram.
CO2
aMDEA
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aMDEA at inlet of absorber
Sweet gas outlet:
15
Sour gas inlet
stripper inlet
CO2 outlet from stripper:
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Regenerated aMDEA
CHAPTER NO.4
4.1 Absorber Design:
Absorber The removal of one or more seleced components from a mixture of gases by absorption into a appropriate solvent. Select between Plate & Packed column:
Vapor liquid mass transfer operation may be carried either in plate or packed column. These two kinds of operation are quitelly different. The relative advantages of plate over packed column are as follows: 1. Plate column are designed to deal wide range of liquid flow rates without flooding. 2. If a system contains solid contents; it will be handled in plate column, because solid will accumulate in the voids, coating the packing materials and making it ineffective. 3. Dispersion difficulties are handld in plate column when flow rate of liquid are low as compared to gases. 4. For large column heights, weight of the packed column is more than plate column. 5. If periodic cleaning is needed, man holes will be provided for cleaning. In packed columns packing must be removed prier cleaning. 6. For non-foaming processes the plate column is preffered. 7. Design information for the plate column is more readily available and more reliable than that for packed column. 8. Inter stage cooling can be provided to remove heat of reaction or solution in plate column. 9. When temperature change is involved, packing may be damaged.
Choice of Plate Type: There are three main types of plate, sieve plate, bubble cap and value plate. We have selected sieve plate because:
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1. 2. 3. 4. 5.
They are lighter in weiht and less expensive. It is easier and cheaper to install. Pressure drop is low as compared to valve and bubble cap plates. Peak efficiency is generaliy high. Maintenance cost is redced due to the ease of cleaning. In case of capacity ratng, sieve plate has high rank as compared to valve and bubble plates.
Sieve plate: Sieve plate is simplest type of cross-flow plate. Vapour passes up through perforations in the plate; and the liquid is retained on the plate by vapour flow. The perforations are usually small holes, but larger holes and slots are used. The arrangement, number and size of the holes are design parameters. Because of their efficincy, wide operating range, ease of maintenance and cost factors, sieve and valve trays have replaced the once highly thought of bubble cap trays in many applications. Factors affecting the absorption Column :
Vapor Flow Conditions: 1. 2. 3. 4.
Foaming Entrainment Weeping/dumping Flooding
Foaming:
Foaming refers to the expansion of liquid due to passage of vapor or gas. Although it provides high interfacial liquid-vapor contact, extreame foaming often leads to liquid buildup on trays. In some cases, foming may be so bad that the foam mixes with liquid on the tray above. Whether foaming will occur depends primarily on physical properties of the liquid mixtures, but is occasionally due to tray desins and condition. Whatever the cause, separation efficiency is always reduced. Entrainment:
Entrainment refers to the liquid carried by vapour up to the tray above and is again caused by high vapor flow rates. It is negative because tray efficiency is reduced: lower volatile material is carried to a plate holdng liquid of higher volatility. It could also contaminate high purity distillate. Excessive entrainent can lead flooding.
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Weeping/Dumping:
This phenomenon is caused by low vapor flow. The pressure exerted by the vapor is insufficient to hold up the liquid on the tray. Therefore, liquid starts to leak through perforations. Excessive weeping will lead to dumping. That is the liquid on all trays will crash (dump) through to the base of the column (via a domino effect) and the column will have to be re-started. Weeping is indicated by a sharp pressure drop in the column and reduced separation efficiency. Flooding:
Flooding is brought about by excessive vapour flow, causing liquid to be entrained in the vapor up the column. The increased pressure from excessive vapor also backs up the liquid in the down comer, causing an increase in liquid holdup on the plate above. Depending on the degree of flooding, the extreme capacity of the column may be severely reduced. Flooding is detecting by sharp increases in column differential pressure and significant decrease in separation efficiency. State of trays & Packing:
Remember that the actual number of trays required for a particular separation duty is determined by the efficiency of the plate. Thus, any factors that cause a decrease in tray efficiency will also change the perfomance of the column. Tray efficiencies are effected by fouling, wear and tear and corrosion and the rates at which these occur depends upon the properties of the liquids being proessed. Thus appropriate materials should be specified for tray construction. Column Diameter:
Vapor flow velocity is dependent on column diameter. Weeping determines the minimum vapor flow required while flooding determines the maximum vapor flow allowed, hence column capacity. Thus, if the column diameter is not sized properly, the column will not perform well Standard Design steps:
1) 2) 3) 4) 5) 6) 7) 8)
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Calculation of theoretical number of stages. Calculation of actual number of stages. Calculation of diameter of column. Calculation of weeping point. Calculation of pressure drop. Downcomer design. Entrainment calculations. Calculation of height of column.
ABSORBER DESIGN
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1.Calculation of theoretical number of stages:35 The main componnt which we want to be absorbed in MDEA is H2S.so, we take it as a reference. H2S: In = 3.73 Out = .0109 Moles of H2S absorbed = 3.719 Eai = = 0.997
or
99.7%
Minimum ⁄ for H2S. (
) min = Ki Eai
L◦ =
lean oil entring absorber.
Vn+1 =
rich gas entering absorber.
Value of K depends on T & P. So, average tower conditions for ki: T = 110 °F P=
= 433psia.
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Ki = 1.6
So, (
)min = 1.6 x 0.997 = 1.5952.
Operating (
)○ = 1.25 (1.5952) = 1.994.
Operating absorption factor Aio = ( 21
)○ .
= = 1.246. Theoretical stages at operating conditions. Eai = AioN+1 – Aio / AioN+1 – 1 0.997 = (1.246)N+1 – 1.246 / (1.246)N+1 – 1 (N+1) log 1.246 = log (
)
(N+ 1)(0.0995) = 1.919 N = 19.09. It means 19 theoretical trays are needed.
1. Calculation of Actual Number of stages: We take 70% efficiency. So, Actual number of stages =
= 27 stages.
Calculation of Diameter of Column:37 Flooding velocity is given by Uf = K1 Where, Uf = Flooding vapor velocity in m/s , base on net column cross-sectional area. K1 = Constant obtained from figure 11.27 vol.6 Coulson & Richardson . FLV = Where, Lw =Liquid mass Flow rate , Vw = Vapour mass Flow rate , In this Case,
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Lw = 148.045 Vw = 17.18
Pv = 21.47
[
]
PL = 1001.48 FLV = = 1.26 We use Plate Spacng 700mm. 38
K1 = .034
Then, UF = 0.034 = 0.23 We take actual velocity as 85% of flooding velocity So, v = 0.85 x 0.23 = 0.20 Maxium volumetric vapor flow rate = =0 .80 Net area required = An = = 4 m2 We take dwncomer area as 12% of total area Column cross sectional area = Ac = = 4.55 m2 Down comr area = Ad = 4.55 -4 = .55m2 Active area , bubbling area = Aa= Ac – 2 Ad = 4.55 – 2(0.55) 23
= 3.45 m2 Total hole aea as 10% of active area , so Hole area = AH = 0.10 x 3.45 = 0.345 m2 Column diameter = Dc = = = 2.40 m.
2. Calculation of Weeping Point: For the calcuation of weeping point, hole diameter must be selected so that at lowest operation rate, the vapor flow velocity is still above weeping point. Maximum liquid flowrate = 148.045 Minimum liquid rate , at 70% turn down = 0.70 x 148.045 = 103.631 x 100 = 39
x 100 = 12%
= 0.77
lw = 0.77 x 2.40 = 1.85m we know how = 750 *
+ 2/3
Lw = weir length, m how = height over weir , mm liquid Lw = liquid flow rate Minimum how = 750 [ = 109mm. We take , hw = 50mm
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] 2/3
hw + how = 109 + 50 = 159mm. 40
K2 = 31.2
Uh(min) = [
]
Uh = minimum vapor velocity through holes, m/s Dh = hole diameter, mm Uh =
]
Uh = 0.77m/s Actual mimum vapor velocity = = = 1.62 m/s.
So, minium operating rate will be well above weeping point.
Calculation of Pressure drop: = 9.81 x 10-3 ht PL = total pressure drop , Pa (N/m2) ht = total pressure drop , mm liquid Total presure drop is giver by ht = hd + (hw + how) + hr ht = total plate presure drop hd = dry plate pressure drop hr = resdual head hw = height of weir how = weir crest, mm liquid hd = 51 [ ]2 Co = Ofice coefficient 25
Uh = Vapor velocity through holes , m/s Uh =
= 2.32 m/s.
We tak carbon steel plate, so plate thickness = 5mm hole diameter = 5mm so, 41
Co = 0.84
hd = 51 [
]2 [
]
= 8.34 mm hr = = 12.8mm ht = 12.48 + 8.34 + 50 + 109 = 179m liquid = 9.81 x 10-3 x 179 x 1001.48 = 1757 Pa = 0.26 Psia (per plate)
Down comer Design: The downomer area and the plate spacing must be such that the level of the liquid and froth in the downomer is well below the top of outlet weir on the plate above. If the liquid rises above the outlet weir the column will flood. hb = (hw + how) + ht+ hdc Where, hb = downcomer bckup, measured from plate surface, mm hdc= head loss in doncomer, mm
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hdc = 166 [
]
where, hdc = head loss in downomer, mm Lwd = liquid flowrate in dwncomer, kg/s Am = Either downcomer area or clearance area under the downcomer Aop which is smaller. Aop = hop Lw Where, hop = height of botom edge of apron above plate Lw = length of weir hop = hw – 10 = 50 – 10 = 40mm So, Aop= 0.040 x 1.85m = 0.074m hdc = 166 [
]2
= 6.62mm So, backup in downcomer = hb= (50 +109) +6.62 + 179 = 34..62mm = 0.33462m Then , backup in downcoer < ½ (plate spacing + weir height) 0.33462 < ½(0.700 + 0.50) 0.33462 < 0.375 So, plate efficiency is acceptable.
Entrainment Calculation: For checking entrainment , we calculate Uv = Uv =
= 0.2 m/s
% flooding =
27
=
= 86%
We already know FLV FLV = 1.26 42
It is well below 0.1, so there is no chance of entrinment and process is satisfactory.
Calculation of Height of Column: No. of plates = 27 Tray spacing = 0.70m Tray thickness = 0.005m Total thickness of trays = 0.135m Top clearance = 1m Bottom clearance = 1m Total height = 20m
4.2 Stripper Design: Stripper: “It is a counter current multi-stage separation column, with liquid feed at top and vapor feed at the bottom stage”. Stripping Phenomenon:
Strippng is a mass transfer operation that involves the transfer of a solute (as H2S & CO2 in our case) from the liquid phase to the gas phase. Stripping Agents:
Air Stream Inert gas Hydocarbon gases Reboled vapors (as in our case)
Types of Stripper:
i.
Refluxed Stripper:
It is emloyed if simple stripping is not sufficient to achieve the desired separation and contacting trays are needed above the feed tray. 28
ii.
Reboiled Stripper:
If the botom product from a stripper is thermally stable, it may be Reboiled at the bottom of the column.
iii.
Open steam/Air stripper: Direct stearm may also be used. Sometimes air or inert gases may also be used (Combination of above can be made based on system’s requirement) Principle of separation: difference in volatilities Created or added phase: vapor Separating agent: stripping vapor
Standard Design Steps:
Calculation of 1) 2) 3) 4) 5) 6) 7) 8)
Theortical number of stages. Actual number of stages. Diamter of column. Weeping point. Pressure drop. Downomer design. Entrainment calculations. Height of column
1. Calculation of theoretical number of stages:43 The main component which we want to be stripped from MDEA is H2S. So, we take it as a refence. Let us supose that 100% of H2S is not stripped and very minute quantities remains in the lean MDEA coming back from Stripper.
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Fraction of H2S stripped = Esi = 0.998 Minimum V/L for H2S = ( )min = Value of K depends on T & P. So, average tower conditions for value of ki: T = 230 °F P = 26psia. 36
Ki = 35
So, ( )min =
= = 0.0285
Operating ( )○ = 1.25 (0.0285) = 0.0356. Operating striping factor Si = ( )○ . Ki = 0.0356 x 35 = 1.246. Theoretical stags at operating conditions. ESi = SiN+1 – Si / SiN+1 – 1 0.998 = (1.246)N+1 – 1.246 / (1.246)N+1 – 1 (N+1) log 1.246 = log (
)
(N+ 1)(0.0955) = 2.0969 N = 20.78 It means 21 theoretcal trays are needed.
2. Calculation of Actual Number of stages: We take 70% efficiency. So,
30
Actual number of stages = = 30 stages.
3. Calculation of diameter of column:37 Flooding velocity is given by Uf = K1 Where, Uf = Floodig vapor velocity in m/s , base on net column cross-sectional area. K1 = Constant obtaind from figure 11.27 vol.6 Coulson & Richardson . FLV = Whre, Lw =Liquid mass Flow rate , Vw = Vapur mass Flow rate , In this case
In this Case, Lw = 315 Vw = 20.38
Pv = 1.96
[
PL = 935.24 FLV = = 0.71 We use Pate Spacing 800mm. 38
K1 = .05
Then, UF = .054 31
]
= 1.8 We take actul velocity as 85% of flooding velocity So, Uv = 0.85 x 1.18 = 1.00 Maximum volmetric vapor flow rate = =10.4 Net area required = An = = 0.4 m2 We take downcomer areaas 12% of total area Column cross sectional are = Ac = = 11.82 m2 Down comer area = Ad = 11.2 – 10.4 = 1.42m2 Active area , bubbling area = Aa= Ac – 2 Ad = 11.82 – 2(1.42) = 8.98 m2 Total hole area as 10% of activ area , so Hole area = AH = 0.10 x 8.98 = 0.898 m2 Column diameter = Dc = = = 3.88 m. Calculation of Weeping Point:
For the calculation of weepig point, hole diameter must be selected so that at lowest operation rate, the vapor flow velcity is still above weeping point. Maximum liquid flowrate = 315 Minimum liquid rate , at 70% turn down = 0.70 x 315 32
= 220.5 x 100 = 39
x 100 = 12%
= 0.77
lw = 0.77 x 3.88 = 2.99m we know how = 750 *
+ 2/3
Lw = weir length, m how = height over weir , mm liquid Lw = liquid flow rate ] 2/3
Minimum how = 750 [ = 138mm. We take , hw = 50mm hw + how = 138 + 50 = 188mm. 40
K2 = 31.2
Uh(min) = [
]
Uh = minimum vapor velocity through holes, m/s Dh = hole diameter, mm Uh =
]
Uh = 6.17m/s Actual minimum vapor velocity = = = 8.11 m/s.
So, minimum operating rate will be well above weeping point. 33
Calculation of Pressure drop:
= 9.81 x 10-3 ht PL = total pressure drop , Pa (N/m2) ht = total pressure drop , mm liquid Total pressure drop is giver by ht = hd + (hw + how) + hr ht = total plate pressure drop hd = dry plate pressure drop hr = residual head hw = height of weir how = weir crest, mm liquid hd = 51 [ ]2 Co = Orifice coefficient Uh = Vapor velocity through holes , m/s Uh =
= 11.6 m/s.
We take carbon steel plate, so plate thickness = 5mm hole diameter = 5mm so, 41
Co = 0.84
hd = 51 [ = 20 mm hr = = 13mm 34
]2 [
]
ht = 188 + 13 + 20 = 221mm liquid = 9.81 x 10-3 x 221 x 935.24 = 2027.6 Pa = 0.29 Psia (per plate)
4. Downcomer Design: The downcomer area and the plate spacing must be such that the level of the liquid and froth in the downcomer is well below the top of outlet weir on the plate above. If the liquid rises above the outlet weir the column will flood. hb = (hw + how) + ht + hdc Where, hb = downcomer backup, measured from plate surface, mm hdc= head loss in downcomer, mm hdc = 166 [
]2
where, hdc = head loss in downcomer, mm Lwd = liquid flowrate in downcomer, kg/s Am = Either downcomer area or clearance area under the downcomer Aop which is smaller. Aop = hop Lw Where, hop = height of bottom edge of apron above plate Lw = length of weir hop = hw – 10 = 50 – 10 = 40mm So, Aop= 0.040 x 2.99m = 0.120m hdc = 166 [
35
]2
= 13.07mm So, backup in downcomer = hb= (50 +138) +13.07 + 221 = 422.07mm = 0.422m Then , backup in downcomer < ½ (plate spacing + weir height) 0.422 < ½(0.800 + 0.50) 0.422 < ½ 0.425 So, plate efficiency is acceptable.
5. Entrainment Calculation: For checking entrainment , we calculate Uv = Uv =
= 1 m/s
% flooding = =
= 85%
We already know FLV FLV = 0.71 42
It is well below 0.1, so there is no chance of entrainment and process is satisfactory. Calculation of Height of Column:
No. of plates = 30 Tray spacing = 0.800m Tray thickness = 0.005m Total thickness of trays = 0.15m Top clearance = 1m Bottom clearance = 1m Total height = 25m
36
4.5 MDEA Surge Drum: Flow rate of MDEA = 2349334.04 Ib/hr Density of MDEA at 160oF = 64.6 Ib/ft3
Basis 24 hours Vol. of MDEA for 24 hr = 2349334.04 x 24/64.6 = 872817.6 ft3 Total vol. of vessel = vol. of MDEA + 10% allowance = 872817.6 x 1.10 = 960099.36 ft3 Let us suppose that 0.3% of MDEA solution is slipped in the surge tank =960099.36 *0.03=2858.29 ft3=81m3 V = d2h/4 Let, h/d = 3 or h = 3d 3 V = 3 d /4 From here, d = 3.25m so, h=9.75 m
[1]
Design Specifications
Time of operation = 24hr Dia. Of vessel = 3.25m Height of Vessel = 9.75m Recommended material of instruction is carbon steel.
37
CHAPTER NO.5 Dynamic Simmulations.
38
CHAPTER NO.6 6.1 Plant Cost Estimation: As the process design is completed it becomes possible to make accurate cost estimation because detailed specification can thus be obtained from various manufactures. However no design project should proceed to the final stages before costs are considers and the cost estimation should be made throughout all the early stages of the design when complete specifications are not available. Evaluation of costs in the preliminary design is said pre design cost estimation. Such estimation should be capable of providing a basis for company management to decide whether or not further capital should be invested in the project. An evaluation of costs in the preliminary design phase is sometimes called as guess estimation and often rule of thumb are used. A plant design obviously must present a process that is capable of operating under conditions which will yield a profit. A capital investment is required to any industrial process, and determination of necessary investment is an important part of plant design project. The total investment for any 39
process consists of physical equipment and facilitates in the plant plus the working capital for money which must be available to pay salaries. Keep raw materials and products on hand and handle other special items requiring a direct cash layout.
6.2 Capital Investment: Before industrial plant can be put into operation, large amount of money must be supplied to purchase and install the necessary machinery and equipment, land services facilitates must be obtained and plant must be erected, complete with all pipe control services. In addition it is necessary to have money available for payment of expenses involved in plant operation. The capital needed to supply the necessary manufacturing and plant facilities is called fixed capital. Fixed cost capital investment while necessary for the operation of the plant termed as Working Capital. The sum of fixed capital investment and the working capital is known as total capital investment. Fixed capital investment classified into two subdivisions: namely
Direct costs Indirect costs
6.2.1
Direct costs:
The direct cost items are incurred in the construction of planet in addition to the cost of equipment:
Purchase equipment Purchase equipment installation Instrumentation Piping Electrical Equipment and materials Building (including services) Service facilities Taxes
6.2.2 Indirect costs: These include:
Design and engineering Contractors expanses Contractors fee Contingency
6.3
Types of Cost Estimation:
Various methods are employed for estimating capital investment are as follows:
40
Preliminary estimate Definitive estimate Detailed estimate
In choosing the method for cost estimation following factors are considered:
Amount of detailed information available Accuracy desired Time spent on estimation
6.4
Methods of Estimating Capital Investment:
Seven methods of estimating capital investment are outlined below:
Detailed item estimate It cost estimate Percentage of delivered equipment cost “Lang” factor for approximation of capital investment Power factor applied to plant capacity ratio Investment cost per capacity Turnover ratio
6.5 Percentage of Delivered Equipment Cost: This method for estimating the fixed or total capital investment requires determination of the delivered equipment cost. The other items included in the total direct plan cost are then estimated as Percentage of Delivered Equipment Cost. The percentage used in making an estimation of this type should be determined on the basis of type of process involved, design complexity required, material of construction, location of the plant, past experiences, and other items depend on the particular unit under consideration.
Purchased equipment cost for common plant equipment Ce= a + b (S)n Where a & b are cost constants S = size parameter n = exponent for that type of equipment
Cost of Absorber:
We know that 41
= Ce = E
C= a+b(S)n Diameter of Absorber = 2.4 m sizing factor (S) = 2.4 46
a = 110, b = 380, n= 1.8 so,
C = 110+380(2.4) = $1947
( Cost/Tray)
Total cost of one absorber = 1947 x 27 = $52569 As we have used two absorber, so Cost of both Absorbers = $52569 x 2 = $105138 This cost is for year 2007, so by applying inflation rate of 3.3% per year, we can find cost in 2013. C = 105138(1.033)6 = $127750.21
1. Cost of Exchanger: We know C= a+b(S)n Sizing parameter of exchanger (S) = 584 m2 a = 1350, b = 180, n = 0.9547
so, C = 1350+180(584)0.95 = $73313 This cost is for year 2007, so by applying inflation rate of 3.3% per year, we can find cost in 2013. 42
C = 73313(1.033)6 = $89080.
2. Cost of Inlet Gas Separator: Diameter = D = 2.06 m Length = L = 6.34 m From Graph48 C = 22000 Pressure Factor = 1.4 (at 30 psia) So, C = 22000 x 1.4 = $30800 This cost is for year 2004, so by applying inflation rate of 3.3% per year, we can find cost in 2013. C = 30800(1.033)9 = $41253.
3. Cost of Lean Solvent pump: Capacity of Pump = 3919 GPM From Graph C = $25000 This cost is for year 1988, so by applying inflation rate of 3.3% per year, we can find cost in 2013. C = 25000(1.033) = $49436.
4. Cost of MDEA Surge Tank: Capacity of tank = 81 m3 We know that C = C= a+b(S)n a = 5000, b = 1400, n = 0.7 (cone roof)49 so,
C = 5000+1400(81)0.7 = $35343
43
This cost is for year 2004, so by applying inflation rate of 3.3% per year, we can find cost in 2013. C = 35343(1.033)9 = $47337. E = $354856
7.6 Direct Cost: Component
% of E
Cost ($)
Purchased Equipment Installation Instrumentation installation Piping Electrical Building Yard improvement Service facilities Land Total direct cost = D = 241262
0.25E 0.07E 0.08E 0.05E 0.05E 0.02E 0.15E 0.01E
88714 24840 28388 17723 17723 7097 53228 3549
7.7 In-Direct cost: Engineering & Supervision
=
0.33E = $117102
(1)
Construction Expenses
=
0.41E = $145490
(2)
Total in-direct Cost
=
$262592
Total Cost
=
direct Cost + indirect cost
=
$503855
Contactor’s Fee
=
X = 0.05 (D+I) = $12063
Contingency
=
Y = 0.10 (D+I) = $24126
Fixed Capital Investment
=
(D+I+X+Y) = $540043
Working Capital Investment =
44
0.15(D+I+X+Y) = $81006
OPERATIONAL PROBLEMS. 7.1
Problems occurring during operation:
One of the reasons that alkanolamine processes have become the predominant choice for both refinery gas giving and natural gas purification is their comparative freedom from operating difficulties. However, several factors can result in undue expense and cause difficulties in the operation of alkanolamine units. Chief among these, from an economic standpoint are corrosion and amine loss. Other operating problems, which occasionally limit the capacity of plant for gas purification, include foaming and plugging of equipment. In many cases, operation can be significantly improved by daily monitoring of key plant operating variables and by proper control and design of treating plant.
7.2
Foaming:
Foaming of alkanolamine solution is perhaps the most common operating problem in amine treating units. It is most frequently encountered in contractor, but may also occur in the stripping column. Causes of Foaming:
Specific cause of foaming includes the following:
45
Water soluble surfactants in the feed gas (e.g. well treating mixes, pipeline corrosion inhibitors) which lowers the amine metaphors surface tension. Excessive antifoam can also cause foaming. Liquid hydrocarbons e.g. entrained compressor lubricating oils in the feed gas or hydrocarbons condensation within the amine absorber. Particulate contaminants (e.g. mill scale, FeS correction products, rust contained in the feed gas or produced within amine treating units. Solids such as FeS do not cause foaming but concentrate at liquid/gas interface and stabilize the foam by increasing the surface viscosity retarding film drainage. Oxygen adulteration of feed gas or amine unit (usually at the amine sump or amine storage tank) and reation of amine heat stable salts. Dissolved iron can catalyze the reaction of amine with oxygen to foam carboxylic acid. Feed gas adulteraction such as carboylic acid, which react with amine to form heat stable salts. Contamination of amine unit with gases and oils during a turnaround. Amines filter elements that have been washed with surfactants or contaminated with oils during manufactue. Contaminants in the amine plant makeup water such as boiler feed water treating chemicals and corrosion inhibitors.
Prevention of Foaming:
Foaming can be reduced or controlled by proper care of the amine solution. The following techniqes reduce the amine solution contamination and minimize foaming:
A properly desined feed gas inlet separator and filter should be provided. A feed gas coalscer should be considered for feed gas stream contaminated with compressor lubricating oils and other finally dispersed aerosols. A properly size slug catcher should be provided if slgs can accumulate in the feed gas line.
A feed gas water wash should be considered when the feed gas streams is severely contaminated wth carboxylic acid or water soluble, surface active pollutes. A feed gas water wash can also remove aerols and ultra fine chemicals.
Onsite of offsie amine solution recovering to remove heat stable salts and amine degradation poducts. No more than 10% of the amine should be tied up as stable salts.
Caustic additin to neutralize heat stable salts to mitigate corrosion and thereby reduce iron sulfite formation.
A properly sized rich amine flash drum remove entrained and dissolved hydrocarbons. Liquid skimming facilitate in the absorber sump, the rich amine flash drum, the regenerator sump and the amine regenerator overhead accumulator.
New plants and old plants that have undergone a major turnaround or often contaminated with oils, greases welding fluxes and corrosion inhibitors. A hot caustic wash (2-5 wt% caustic soda) followed by a hot condensate wash can remove these impurities and help to prevent foaming.
Addition of antifoam is carried out. 7.3 Corrosion: By far the most serious operating problem encountred with amine gas purification process is corrosion as would be expected this problem has been given widest attention. Generally, it occurs in regenerator heat exchanger and pumps. The extent and type of corrosion has been observed to depend upon such factors as the amine used, the presence of impurities in the solution leading with acid gas, the temperature and pressure, predominant in various part of the plant, the velocity with which the solution flows and others. However, it appears that the principal corroding agents are the add gases. The rate of corrosion growths with increase acid gas concentration n solution. Corrosion due to hydrogen sulfide and carbon dioxide is frequently observed a filter shell and the hot end heat exchanger tubes. To minimize corrosion by hydrogen sulfide and carbon dioxide, the acid gases must be shell in a relatively corrosive form until regeneration of amine solution is stripping still Overloading the amine solution will increasethe casual for corrosion due to pressure discount or high temperature in the heat exchanger. This danger can be remedied bymaintaining adequate pressure on the amine solution and by operating the unit at as low and acid gas alkanolamine ratio as possible. This ration should not exceed 0.05
46
moles of acid gas per mole of alkanolamne and should be event less of condition licences Mechanism of Corrosion:
It is known that free or “aggressive” carbon dioxide causes severe corrosion particularly at elevated temperature and in the presence of water It is believed that the metallic iron with carbondi acid which results in the formation of stable iron bicarbonate. Further heating of solution any cause the release of carbon dioxide and the precipitate of the iron as the relatively insolule carbonate. Hydrogen sulfide attacks steel as an acid with the subsequent formation of insoluble ferrous sulfite. This compound forms a coating on the metal surface which does not adhere tightly and therefore affords little protection from further corrosion. There is no satisfactory correlation available for carbon dioxid hydrogen sulfide mixture, which relates the corrosive attacks to be probable with any givenratio of hydrogen sulfide to sulfur dioxide. However, certain generalized obseration has been made. It appears that in plant handling predominantly carbon dioxide, ver small extent of hydrogen sulfide may actually reduce corrosion. On the other hand, eac of the acid gases growths the corrosive attacks of the other Methods of Minimizing Corrosive Attacks:
Corosion can be reduced by various methods, including certain protection in the operation ad process design of purification plants. Use of more expensive corrosion resistant material nd continues or periodic removal of corrosion promoting agent from the solution. Acombination of several of these measures usually leads to most satisfactory and e\conomical to reduce corrosion attacks;
The temperature of the solution in the reboiler and the temperature of the steam usedin the reboiler should be kept as low as proble. Use of high temperature heat carrying media, uch as oil, should be avoided to maintain the lowest possible skin temperature of metal Pressure regenerator with its supplemenry high temperatures results in severe corrosion of reboiler tubes; it is, therefore, good ractice to maintain the lowest possible pressure on the stripping column and reboiler To prevent oxygen from entering te system, it is prudent to maintain a blanket of inlet gas over all serving of the solutin, which could be exposed to atmosphere and to ensure the pressure the suction side ofall pumps. Continuous removal of suspeded solids (by nitration) and the decomposition product (by distillation of a side stream) enerally helps to reduce corrosion.
7.4
Chemical Losses
The loss of a solvent can bea serious operating difficulty in alkanolamine gas purification plants. Corrosion can be icurred by entrainment of the solution in the gas stream 47
vaporization or chemical dgradation of the amine. Loss of the solvent by entrainment or vaporization is undesirable \not only because of the cost of chemicals but also because of the contamination of the ipelines by liquid deposited on the walls. In addition when alkanolamine solution are ued to purify the gas to be used in catalytic process, entrainment by vaporization of solvet esult in a serious poisoning of the catalyst.
7.5
Losses due to Volatility:
Glycol volatility losses are usually significant in ethylene glycol, di-ethylene glycol but very less in tri-ethylne and higher glycols, which have very high boiling points. Hence usually a very small amunt of glycol is lost by evaporation into gas stream in absorbers and also in regeneators.
Prevention from Volatility Losses: Volatility losses can be prevented by following methods:
7.6
A cold water flux is convyed at the top plate of regenerating column. Normally absorbers are operted at lower temperatures (80-1100F recommended) to avoid losses.
Entrainment:
In many cases most of the glycol loss occurs as carry over of solution with the product gas. Entrainment losses are fashined either by inefficient mist withdrawal or by foaming and subsequent carry over solution. Entrainment losses from glycol absorber vary significantly contingent on the mechanical design of both the upper solution of absorber and mist elimiation devices.
Prevention from Entrainment Losses: Entrainment can be mi diminishd by the following techniques:
Using efficient mist eliminaton equipment. Application of the foam inhibitor.
References:
1.
ARTHUR KOHL AND RICHARD NELSON, “Gas Purification”, Edition 5th.
2.
Basic Principles and Calculations in Chemical Engineering.7th Edition.
3.
Coulson & Richardson's Chemical Engineering - Volume 6.
4.
GAS PROCESSORS SUPPLIERS ASSCIATION,” Engineering Data Book
48
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