FVO Nr.
FabnllsvooDOntwerp
Vakgroep Che..sche Procestechnologie
Subject
The producti of tluosulphuric acid out of
sulphur tritllile and hydrogen fluoride
Authors
L.P.A.F. EIst
K.B.
Geerse
M.D.
Jager
M.H.
Paans
Keywords
Phone
015.2613296
015.2573262
015.2567354
015.2130148
Sulphur burning, ~ur dioxide, Sulphur trioxide, Hydrogen fluoride,~c distillation, Fluosulphuric acid
Date of assigment
Date of report
::
March 1996
lune 1996
FVO Nr.
Fabrieksvoorontwerp
Vakgroep Chemische Procestechnologie
Subject
The production of fluosulphuric acid out of
sulphur trioxide and hydrogen fluoride
Authors
L.P.A.F. EIst
K.B.
Geerse
M.D. Jager
M.H. Paans
Keywords
Phone
015.2613296
015.2573262
015.2567354
015.2130148
Sulphur burning, Sulphur dioxide, Sulphur trioxide, Hydrogen fluoride,Azeotropic distillation, Fluosulphuric acid
Date of assigment:
Date of report
March 1996
June 1996
Summary
Fluosulphuric acid (FSA) is used as a catalyst in the production of Terathane® at Ou Pont de Nemours in Dordrecht. This catalyst is nowadays bought from Bayer in Ludwigshafen and transported by train to Dordrecht. The severe acidity and toxicity of F8A gives the transport an extra risk. Therefore, DuPont would Iike to produce FSA on site. The total demand for FSA is 11000 ton per year.
The reactants, used for the production of F8A, are sulphur trioxide (803) and hydrogen fluoride (HF). The design is divided in three parts. The first part is the production of pure HF, the second part is the production of 803 and the third part is the production of F8A.
Part one.
HF is produced out of an aqueous waste stream from the Freon-22® plant on site. This waste s1ream contains 15 wt% HF, which can be stripped by use of sulphuric acid. The proposed design yields a product s1ream of 2208 ton per year pure HF, which is enough for 1he FSA plant. Out of an economic point of view, this process is not feasible. Due to high investments and manufacturing costs, the recycling of HF not a paying process.
Part two.
803 is produced out of sulphur, which is combusted to S02 and th en converted to S03 by use ofV20s as catalyst. A quantity of 4280. ton per year of liquid sulphur is converted to 8850 ton per year pure S03. The main part of the S03 is condensed and sent to the FSA unit. The rest of the S03 is scrubbed with a sulphuric acid stream to concentrate it. This sulphuric acid stream is used to dry the feed air stream. Financially, this is a paying process.
Part three.
The production of FSA is done by adding liquid HF and liquid S03 to a recycling FSA stream. The heat of reaction is removed by using an air cooler. The produced FSA is split oft the reCJcle stream and sent to a storage tank, after which it is transported to the Tera1hane plant. This part of the plant is profitable .
~. mancIa overvIew . I 0 fth
e
olan. It
Plantwith Plant without
HF-regeneration HF-regeneration
Totallnvestment (Mfl) 54.4 36.5
Table of contents
1.1ntroduction... 1
2. Part one: HF concentration... 2
2.1. Basis of design... 2 2.1.1. Feed ... 2 2.1.2. Product specifications... 2 2.1.3. Utilities... .... 2 2.1.4. List of chemicais... 3 2.1.5. Literature... ... ... 3 2.2. Process description. ... ... ... 4
2.2.1. General process structure... 4
2.2.2. Description of sections... 4
2.2.3. Mass balance... ... 5
2.2.4. Simulating in ChemCAD... 5
2.2.5. Literature... ... 5
2.3. Equipment design... ... .... ... ... 6
2.3.1. Design of a distillation column 6 2.3.2. Design of a heat exchanger ... 10
2.3.3. Literature ... 12
2.4. Nomenclature. ... ... ... 13
3. Part two :Sulphur combustion ... 14
3.1. Basis of design... ... 14 3.1.1. Feed ... 14 3.1.2. Product specifications ... 14 3.1.3. Utilities... ... 14 3.1.4. List of chemicais... . ... 14 3.2. Process description... ... 16
3.2.1. General process structure ... 16
3.2.2. Description of sections ... 16
3.2.3. Mass balance ... 18
3.2.4. Simulating in ChemCAD ... 19
3.3. Equipment design ... 20
3.3.1. Air dryer ... 20
3.3.2. Sulphur combustion reactor ... 21
3.3.3. Sulphur dioxide converter ... 22
3.3.4. Sulphur trioxide absorber ... 23
3.3.5. Sulphur dioxide absorber ... 23
3.3.6. Nomenclature ... 24
3.4. Literature ... 25
4. Part three: Conversion to Fluosulphuric acid ... 26
4.1. Basis of design ... 26 4.1.1. Feed ... : ... 26 4.1.2. Product specifications ... 26 4.1.3. Utilities ... 27 4.1.4. List of chemicais ... 27 4.1.5. Literature ... 27
4.2. Process description... 28
4.2.1. General process structure ... 28
4.2.2. Oescription of sections ... 28
4.2.3. Mass balance ... 29
4.2.4. Simulating in ChemCAD ... 30
4.2.5. Literature ... 30
4.3. Equipment design ... 31
4.3.1. Air cooler design ... 31
4.3.2. Reactor design ... 35
4.3.3. Literature ... 35
4.4. Nomenclature ... 36
5. Health, safety and environment.. ... 37
5.1. Chemicais ... 37 5.1.1. Hydrogen fluoride ... ~ ... 37 5.1.2. Sulphur ... 37 5.1.3. Sulphur dioxide ... 37 5.1.4. Sulphur trioxide ... 38 5.1.5. Sulphuric acid ... 38 5.1.6. Fluosulphuric acid ... 39 5.2. HAZOP analysis ... 40
5.2.1. Oow fire & explosion index.... 40
5.2.2. Sulphur combustion unit.. ... 41
5.2.3. Sulphur trioxide reactor ... 42
5.2.4. Sulphuric acid distillation ... 43
5.2.5. FSA reactor ... 44
5.3. Treatment of waste streams ... 45
5.3.1. Waste water ... 45
5.3.2. Oft-gas treatment.. ... 45
5.3.3. Literature ... 45
6. Process Control.. ... 46
6.1. Product quality control ... 46
6.2. Plant startup and shutdown ... 48
6.3. Literature ... 49 7. Cost Engineering ... 50 7.1. Introduction.... ... ... 50 7.2. Investment calculation ... 51 7.2.1. Introduction ... 51 7.2.2. Taylor's method ... 51 7.2.3. Zevnik-Buchanan's method ... 51 7.2.4. Investment calculations ... 52 7.3. Calculation of costs ... 53
7.3.1. Volume dependent costs ... 53
7.3.2. Cost of labour ... 53
7.3.3. Total costs ... 54
7.6. Profitability analysis ... 57
7.6.1. Return on investment.. ... 57
7.6.2. Internal rate of return ... 57
7.6.3. Pay out time ... 58
7.7. Literature ... 59
7.8. Nomenclature ... 60
8. Conclusions and recommendations ... 61
Appendix A: Numeric results ... .. Calculation of HF column ... A 1 Results HF and H2S04 columns ... A6 Results of air cooler ... A7 Appendix B: Numeric economie results ... .. Taylor's method including HF unit.. ... B1 Taylor's method excluding HF unit.. ... 83
Zevnik-Buchanan's method... 85
Appendix C: Streamcompositions ... . Streamcompositions of S03 unit.. ... C 1 Streamcompositions of HF & FSA unit.. ... C2 Mass & heat balance S03 unit.. ... C3 Mass & heat balance HF unit.. ... C6 Mass & heat balance FSA unit... C7 Appendix 0: Equipment list.. ... . Heat exchangers ... 01
Pumps, compressors and expanders... 03 Appendix E: Specification sheets ... .
Heat exchangers ... E1 Columns ... E6 Reactors... E 11 Appendix F: Flowsheets ... .
Flowsheet of S03 unit.. ... F1 Flowsheet of HF & FSA unit.. ... F2 Appendix G: Patents ... .
Patent used in HF unit.. ... G 1 Patent used in FSA unit.. ... G2
1. Introduction
This report is written as a result of a design study made for DuPont de Nemours Netherlands. The objective is to investigate by which process on site production of the chemical fluosulphuric acid (FSA) in Dordrecht is possible and economically feasible. Fluosulphuric acid is a very streng acid, which is very corrosive in contact with water. Fluosulphuric acid is used to catalyse the polymerisation of
tetrahydrofuran to polytetramethylene-ether-glycol. The maximum consumption of fluosulphuric acid is 30 tons per day. The design capacity of the plant is assumed to be 11000 tons per year.
At this moment the only producer of fluosulphuric acid in Europe is Bayer. So
fluosulphuric acid is purchased by DuPont at Ludwigshafen, from where the chemical is transported by train to Dordrecht. The price of fluosulphuric acid is f 1650,- per ton. DuPont does produce fluosulphuric acid in the United States, in La Porte (Texas). Shipping fluosulphuric acid from Texas to Europe is impossible, because safety measures for handling fluosulphuric acid are very stringent.
As part of the study, the economic feasibility of production of anhydrous hydrogen fluoride from a waste stream containing 15 wt% hydrogen fluoride in water, which is available on site, is investigated.
Fluosulphuric acid is produced by reacting hydrogen fluoride (HF) with sulphur trioxide (S03). The process needed for th is reaction is fairly simpie. The problem is how to come at the reactants. Both reactants react violently with water and are
difficult to handle. Furthermore, most materials are not resistant to hydrogen fluoride, especially in contact with water. Basically, there are two ways to come at both
reactants:
803: 1. Producing sulphur trioxide by buming sulphur with air or pure oxygen and converting the sulphur dioxide by use of a catalyst.
2. Producing sulphur trioxide by distillation from oleum.
HF: 1. Producing hydrogen fluoride from the aqueous waste stream available on site in Dordrecht.
2. Simply using pure hydrogen fluoride, bought from a supplier.
In this report production of sulphur trioxide from sulphur and air and production of anhydrous hydrogen fluoride from a 15 wt% aqueous solution is discussed. The report is divided in three parts: one part discusses the production of hydrogen fluoride, a second part discusses the production of sulphur trioxide, a third part discusses the production of fluosulphuric acid. Because all three chemicals are inherently dangerous a more general analysis of the safety and the control of the process is inc1uded. An estimation of the total costs of a fluosulphuric acid plant is also made.
In short, the objective of the report, designing a process capable of producing 11000 tons per year of fluosulphuric acid in an economical way, is met as follows:
- Giving details of the preposed process, taking into account the situation and availability of utilities at the site in Dordrecht.
2. Part one: HF concentration
2.1. Basis of Design
In this chapter the basis of design of the purification of a hydrofluoric acid (HF) wastestream is discussed. Feed and product specifications are given as weil as a specification of utilities and a list of used chemicais.
2.1.1. Feed
The hydrofluoric acid feedstream is a wastestream of the fluorcarbon-production. This wastestream has to be concentrated before it enters the fluosulphuric acid plant.
The following assumptions are made conceming availability and composition of the feed:
1. the feed consists of a continuous flow of 15 wt% HF in water.
2. the massflow of the feed is 1870 kg/ho This massflow rate is just large enough to produce the needed pure HF for the production of FSA.
3. the feed contains some high-boiling organic impurities, which are assumed to be neglectable.
4. specifications of the feedstream: - temperature - pressure - phase
2.1.2. Product specifications
25°C 1 bar liquidThe productstream is substantially pure HF (99.8 wt%,liquid), at 10°C. The mass flow rate of the product is 276.0 kg/ho
2.1.3. Utilities
The following utilities are available on site [1]:
- Steam.
Superheated steam is available on site at two different pressures. T bi 211 P
a
e
.
.
.
.
rope les rf 0 f supe eae seam rh t d t .Pressure Processtemperature Condensationtemperature
(bar) (OC) (OC)
10 220 180
40 410 250
- Electricity.
Three voltages are available, 220, 380 and 10.000 VAC.
-Air.
Process air, 20°C, 7 bar, is available. Dewpoint of the air is 40°C, the relative humidity is 70 %.
-Water.
a. Drinking water is available at 7 bar.
b. Demineralized water is available at 7 bar.
c. Cooling water is available at 3 bar, ground level. The inlet-temperature is 20°C, the maximum outlet-temperature is 40°C.
2.1.4. List of chemicals
The following table contains the chemicais, used in the HF concentrating process. Table 2.1.2: List of chemicals.l21l3j.
Hydrogenfluoride Water HF H20 boHing point 19.5 100.0 (OC) melting point -83.0 0.0 (OC) densitr 958.0 1000.0 (kg/m ') molemass 20.0 18.0 (kq/kmole) MAC 2.5
-(ppm) lethal dose 774.0 -(ppm) prize 0.00 0.00 (Dfllton)2.1.5. Literature
(1]
J
.
Grievink, C.P. Luteijn and M.E.A.M. Thijs-Krijnen, 'Handleidingfabrieksvoorontwerp', TU Delft, 1994. Sulphuric acid (98 wt%) H2S04 279.6 10.4 1843.0 98.1 1.0 135.0 130.0
2.2. Process description
This chapter gives a complete description of the HF-clean-up process as weil as an overall mass balance of the process.
2.2.1. General process structure
The process of concentrating the aqueous HF stream is given in the following figure. The flows are given in kg/ho
'"
H2SOiH2Owaste HF to
water 1084 storage 275
J..
3058199I
o
feed HF distillation HF stripsectien H2SO.ol 2801159~Jio.
r+
r+
distillatien,.
0
HFIH20
I
H2SO./H2O lwaste water2751511 3058/610 511
Figure 2.2.1: Process structure of HF concentration. The different sections of the process are discussed below.
2.2.2.
Description of sections
-HF distillation
The first step in concentrating the HF, is the distillation of the feed to its near-azeotropic composition.
The feed, 15 wt% HF, enters the first distillation column at a mass flow rate equal to 1870 kg/ho The feed is at atmospheric pressure and temperature (1 bar, 20°C, liquid phase). The feed is concentrated to a stream containing 35 wt% HF, i.e. 1084 kg/h of wastewater (1 bar, 40°C) is removed. This waste stream contains 0.5 wt% HF (pH 2.2), that will be neutralised by adding calciumcarbonate. The deposition is removed afterwards.
-HF stripsection
The concentration of HF is not the key-issue of our design, therefore the HF
stripsection is treated as a black box model. A complete description can be found in [1], appendix G. The only concerns of the design are the incoming and outgoing streams. A heat balance is also made, to calculate the energy costs.
A sulphuric acid stream is used to dehydrolyse the HF stream. Contrary to [1], the used sulphuric acid is recycled after distillation, instead of using the acid once-through.
A strearn of 786 kg/h near-azeotropic HF enters the stripsection and is being stripped Of water by adding a stream of 3157 kg/h 98 wt% H2
S0
4 . This H2S0
4 leaves the black box as a 83 wt% H2S0
4 stream. The process is operated at 1 bar and 120°C. HF is formed in the vapour phase, liquefied (1 bar, 10°C) and subsequently stored. The HF stream is 276.0 kg/h, with 0.2 wt% H20.HydrogenFluoride / Water at 1.01 bar By NRTL Temp (Cl 120 ...: ...: ...: ... :
--
... , - - - , · . · . 110 : --.--_···· __ ···_··1-·~, ··-~···_···:··_, ··· __ ··· ... .. ... 100.
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: Job Name: Case Name: 06-24-96 HF-H20 HF-H20 11: 15 .... : : : ... ... : :1I
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Temp (Cl 400 i I ! 0.2 0.4 0.6 0.8 Xl 1 Vl Weight FracSulfuric Acid I Wate~ at 1.01 bar 8y NRTL
~ . ~
-1.0 Job Name: H2S04H20 H2S04H20 11: 24 350 ... :.:.--Case Name: ... ~";I 06-24-96~
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-H2S04 distillation
Before recycling the spent sulphuric acid to the stripsection, it is distilled to near-azeotropic composition.
The feed (1 bar, 120°C, Iiquid, 3668 kg/h), coming from the stripsection, contains about 83 wt% H2S04 and is distilled to 97 wt% H2S04 (1 bar, 365°C, liquid phase),
i.e. a stream of 511 kg/h wastewater (1 bar, 40°C) is removed. The H2S0
4-concentration in the wastewater (2 ppb, pH 6.8) is neglected.
The lean H2S04 -stream (3157 kglh) is redirected to the stripsection.
2.2.3. Mass balance
In figure 2.2.2, the overall mass input and output is shown. The sulphuric acid stream is recycled, so there is only need for a sm all input of fresh sulphuric acid, to keep the stream at a constant volume.
H 2
S04
waste HF feedProcess HFw aste HF pr oduct
Figure 2.2.2: Overall mass input and output.
T bi 221
a
e. .
.
.
1
n- an oU[pu s reams d t t t 0 f th e process.HF H20 H2S04 (kglh) (kg/h) (kglh) HF feed 280.0 1590.0 0.0 H2S04 waste 0.0 511.0 neglectable HF waste 5.0 1079.0 0.0 HF product 275.0 0.0 0.0
2.2.4. Simulating in ChemCAO©
Both sulphuric acid/water and hydrofluoric acid/water systems form an azeotrope, at 98 wt% and 37 wt% respectively. Also both systems are electrolytes, so at high concentrations of H2S04 and HF activity coefficients have to be used. To simulate
the process in ChémCAD [2] the NRTL-electrolyte-model is chosen to describe the behaviour of the acid/water systems. Comparison of literature [3] and the ChemCAD-data, shows that the ChemCAD model is satisfying.
Latent heat is chosen as enthalpy model.
The stripsection is considered to be a black box, therefore the H2SOJHF/water
system is not modelled.
2.2.5. Literature
[1] M. Provoost and
J
.
Chastel, 'Preparation of anhydrous hydrofluoric acid',US Patent Office, 2,952,334 (1959) [2] ChemCAD
111,
'user's manual[3]
J.
Gmehling et al. , 'Vapor-Iiquid equilibrium data coll., DECHEMA 1988 52.3. Equipment design
In this chapter the design of distillation columns and heat exchangers is discussed. Specific calculations are included in appendix A.
2.3.1. Design of distillation column
In the process, two distillation sieve-tray columns are used; one to distillate HF and the other one to distillate H2S04. Both distillations are non-ideal, therefore Frenske's method of calculating the number of theoretical plates cannot be used. The McCabe-Thiele graphical method is used instead.
For the design of the whole column, the method as discussed in [1], is used.
-Number of actual plates
With an assumed column efficiency of 75 per cent, the number of actual plates is given by:
-Column sizing
Ntlleo
N<lCt=--0.75 (2.3.1)
The flooding condition fixes the upper limit of vapour velocity. A high vapour velocity is needed tor high plate efficiencies, and the velocity wil! normally be between 70 to 90 per cent of that which would cause tlooding. For design, a value of 85 per cent of the flooding velocity is used.
The flooding velocity is estimated by:
!
P
L -Py
uf = KI'
I
(2.3.2.)~
Py
Kl
is obtained trom figure 11.29 [1]. The liquid-vapour flow factor, FLv is given by:
r-F =
L
w.
J
P
Y
LV
V
P
w L
(2.3.3.)
For liquids with other surface tensions than 0.02 Nim, correction for
Kl
is needed:(
cr
)
0
.
2
K = K ·-l,c . I 0.02 (2.3.4.)
To calculate the column diameter an estimation of the net area An is required. As a first trial the downcomer area is taken as 12 per cent of the total, and the ratio hole area lactive area is assumed to be 10 per cent.
-Liquid flow
The choice of plate type (reverse, single pass or multiple pass) will depend on the liquid flow-rate and column diameter. An initial selection can be made using eoulson and Richardson,[1].
-Entrainment
Entrainment can be estimated from the correlation given in Fig 11.29 [1], which gives the fractional entrainment 'I' as a function of the liquid-vapour factor FLv, with the percentage approach to flooding as a parameter.
The percentage flooding is given by:
-Weep point u %flooding= ~ u f (2.3.5.)
The lower limit of the operating range occurs when liquid leakage through the plate holes becomes excessive. This is known as the weep point. The vapour velocity at the weep point is the minimum value tor stabie operation. The hole area must be chosen so that at the lowest operating rate the vapour flow velocity is still weil above the weep point.
The minimum design vapour velocity is given by:
(2.3.6.)
-Weir liquid crest
The height of the liquid crest over the weir can be estimated using the Francis weir formula, which can be written for a segmental downcomer as:
2
h
:=750
.(
~
)
3
ow P .1 L w (2.3.7.) -Hole pitchThe hole pitch lp should not be less than 2.0 hole diameters, and the normal range will be 2.5 to 4.0 diameters. Within this range the pitch can be selected to give the number of active holes required for the total hole area specified.
Square and equilateral triangular pattems are used; triangular is preferred. The total hole area as a fraction of the perforated area Ap is given by the following expression, for an equilateral triangular pitch:
(2.3.8.)
-Plate pressure drop
The pressure drop over the plates is an important design consideration. There are two main sourees of pressure 1055: due to vapour flow through the hole and due to the statie head of liquid on the plate.
A simple additive model is normally used to predict the total pressure drop. The total is taken as the sum of the pressure drop calculated for the flow of vapour through the dry plate (the dry plate drop hd), the head of clear liquid on the plate (hw
+
how), and a term to account for other, minor, sourees of pressure loss, the so-called residualloss hr. The residualloss is the difference between the observed experimental pressure drop and the simple sum of the dry-plate drop and the clear-liquid height. It accounts for two effects: the energy to form the vapour bubbles and the fact that on anoperating plate the liquid head will not be clear liquid but a head of "aerated" liquid froth, and the froth density and height will be different from that of the clear liquid. It is convenient to express the pressure drops in terms of millimetres of liquid. In pressure units:
(2.3.9.)
Dry
plate dropThe pressure drop through the dry plate can be estimated,using the following equation:
(2.3.10.)
Where the orifice coefficient Co is a function of the plate thickness, hole diameter, and the hole to perforated area ratio. Co can be obtained from Fig. 11.34.[1] Residual head
The following equation is used to estimate the residual head:
Total plate pressure drop
12.5.103
h :=
-r
The total plate pressure drop can now be written as follows: h t := h d +- (h w +- h ow ) +- h r
-Downcomer design (back-up)
(2.3.11.)
(2.3.12.)
The downcomer area and plate spacing must be such that the level of the liquid and froth in the downcomer is weil below the top of the outlet weir on the plate above. If the level rises above the outlet weir the column will flood.
The back-up of liquid in the downcomer is caused by the pressure drop over the plate (the downcomer in effect forms one leg of a U-tube) and the resistance to flow in the downcomer itself. Therefore the downcomer back-up is given by:
The he ad loss in the downcomer can be estimated using the equation below:
h :=
166.(~)2
de
Pr..A
m
(2.3.14.)
The clearance area under the downcomer is given by: A := h ·1
ap ap w (2.3.15.)
Froth height
To predict the height of "aerated" liquid on the plate, and the height of froth in the downcorner, some means of estimating the froth density is required. For design purposes it is satisfactory to assume an average value of 0.5 of the liquid density. This value is also taken as the mean density of the fluid in the downcomer; which means that for safe design the clear Iiquid back-up, calculated from equation (2.3.13.), should not exceed half the plate spacing, It, to avoid flooding:
(2.3.16.)
-Downcomer residence time
Sufficient residence time must be allowed in the downcomer for the entrained vapour to disengage from the liquid stream; to prevent heavily "aerated" liquid being carried under the downcomer. A time of at least 3 seconds is recommended for design purposes.
The downcomer residence time is given by: A·h d 'be L 'p
t .
-r
9
2.3.2. Design of heat exchanger
In the HF process four heat exchangers are used, two to cool down the wastewater streams from the distillation columns (H37 and H43), one to cool down the sulphuric acid stream (H44) and one to cool down the HF product stream (C47). The common design method that is used for acondenser with water, is taken from [1]. Results for the heat exchangers can be found in appendix E.
First choose the type of heat exchanger/condenser, i.e. horizontal or vertical type, number of tube passes, U-tube or floating head. Also must be considered which flow goes through the tubes and which one through the shell. When those evaluations are made, the design can start.
-Area required
The area required for the heat exchanger depends on the heat duty, the corrected temperature difference and the heat-transfer coefficient.
H A :=
-U·~Tm (2.3.18)
The corrected temperature difference is derived from the logarithmic temperature difference, which is given by:
(2.3.19)
This temperature difference is corrected by a factor Ft, which is graphically determined in figure 12.19 [1]. This factor depends on the parameters Rand S, which are given by:
Tl - T 2 R := -t 2 - tI S:= -T 1 - ti (2.3.20) (2.3.21)
The overall heat-transfer coefficient is estimated to determine the required area. This estimation is later on compared to the calculated value and, if necessary, changed by an iterative calculation.
Atube is calculated with the outside tube diameter. The pitch is 1.25 times this diameter.
The tube bundie diameter is now given by:
1
Db=dor~u~f
(2.3.23)K1 and n1 are parameters, depending on the pitch and the number of tube passes. The values can be found in table 12.4 [1].
When the condensing stream flows through the shell side, which is usual, a wall temperature of the tubes has to be calculated. This temperature depends on the condensing coefficient and the mean shell and tube temperature:
-Condensing coefficient Tl"" T 2 T := -2 tI + t 2 t := -2 (2.3.24) (2.3.25)
The condensing coefficient has to be estimated, like the overal! coefficient. The wal! temperature is then also an estimation of the real wal! temperature:
(2.3.26)
With the estimated wal! temperature, the physical properties of the condensate are calculated, and with those, the real condensing coefficient.
Wc
r
h:= -L-N tubes 1 1 (2.3.27) (2.3.28)If the estimation of the condensing coefficient is the same as the one in 2.3.26, the wal! temperature is correct. Else, a new estimation has to be made.
- Tube side coefficient
The tube side coefficient, in this case with water in the tubes, can be calculated with
this simplified expression for water:
-Overall coefficient 0.8 Ut h i := 4200·( 1.35 + 0.02·t)·-d .0.2 1 (2.3.29)
The only parameters, that are obligatory to calculate the overall coefficient, are the Fouling factors, taken from tabla 12.2 [1], and the thermal conductivity of the tube wall material, kw, taken from table 12.6 [1]. The overall coefficient is given by:
d
o·ln
(
~dol
.)
1 do 1 do 1
:= - +-- + +-_ . - +- _ . - (2.3.30)
U he h od 2· k w d i h id d i h i
If the overall coefficient has the same value as the estimated one at the beginning of the design, the calculation is complated. Else, a new estimation has to be done, and the calculation must be repeated.
-Pressure drop
There are two pressure drops in the exchanger, one in the tubes and one in the shell. The pressure drop in the tubes is given by:
(2.3.31 )
The pressure drop of the shell side is given by:
(? "
_
.
.).
. ) -"?)2.3.3.
Literature
[1] eoulson & Richardson, 'Chemical Engineering, Volume 6, design', Pergamon
2.4. Nomenclature
Distillation column design Heat exchanger design
Aap Clearance area under apron m2 A Heat transfer area m2
Ao Downcomer area m2 Alube Tube area m2
~ Total hole area m2 Db BundIe diameter m
Am
Same as AJAp m2 Os Shell diameter mAp Perforated area m2 de Equivalent diameter m Co Orifice coefficient dj Inside tube diameter m dh Hole diameter m do Outside tube diameter m
FLV Column liquid-vapor factor FI Temperature correction factor m hap Apron clearance m 9 Gravitational acceleration ms02 hb Height of Iiquid backed-up m H Total heat to exchange J
hbc Oowncomer back-up m he Heat transfer of condensate Wm02K"' hd Dry plate pressure drop m hj Heat transfer of tube inside Wm02K"' hdc Head loss in downcomer m hid Fouling coefficient inside tube Wm02K"' how Height of Iiquid crest m hod Fouling coefficient outside tube Wm02K"' hr Plate pressure drop m jt Friction factor
hl Total pressure drop m K, Constant from table 12.4 [1]
hw Weir height m kt Conductivity of Iiquid Wmo'K"' K, Constant used in 203.4 kw Conductivity of tube wall Wmo'K"' K2 Constant used in 20306 L Tube length m
lp Pitch of holes m Ib Baffle spacing m
II Plate spacing m Ntubes Number of tubes
Iw Weir length m .1Ps Pressure drop in shell Pa Lw Liquid mass flow-rate kgsO' .1Pt Pressure drop in tubes Pa lwd Lw through downcomer kgsO' PI tube pitch m Nact Actual number of plates R Temperature ratio
N1tIeo Theoretical number of plates S Temperature ratio
tr Residence time in downcomer s T, Shell inlet temperature K Ut Vapor velocity at flooding point mso' T2 Shell outlet temperature K
Uh Vapor velocity through holes mso' T Shell temperature K Un Net vapor velocity ms
o
' t, Tube inlet temperature K
Vw Vapor mass flow-rate kgsO' t2 Tube outlet temperature K
PL Density of liquid kgm03 t Tube temperature K
pv Density of vapor kgm03 óT1m Logarithmic mean temperature K
cr Surface tension Nmo' óTm Corrected mean temperature K
tw Wall temperature K
U Overall heat transfer coefto Wm02Ko'
Us Shell side velocity mso'
Ut Tube side velocity ms 0'
Wc Condensate mass-flow kgsO'
~ Viscosity of condensate Paos Pt Liquid viscosity kgm03
pv Vapor viscosity kgm03
r
h Condensate loading kgmO'sO'3. Part Two: Sulphur Combustion
3.1. Basis of Design
In this chapter the basis of design of the oxidation of sulphur tosulphur dioxide and sulphur trioxide is discussed.
3.1.1. Feed
The feed is a flow of liquid sulphur which will be combusted with air. The air must be dried before it can be used in the process, because the water in the air will react with the produced sulphur trioxide to sulphuric acid. The sulphuric acid will end up in the sulphur trioxide product stream.
The fol/owing assumptions are made conceming availability and composition of the feed:
1. The liquid sulphur will be provided at a temperature of 140°C. 2. The impurities in the sulphur are assumed to be neglectable. 3. The relative humidity of air is assumed to be 70 %.
4. The composition of the dry air is 23.2 wt% O2 and 76.8 wt% N2 5. Specifications of the air feedstream: - temperature 25°C
- pressure 1 bar
3.1.2. Product specifications
The main product is substantially pure sulphur trioxide (99.8 wt%, liquid, 30°C) with a trace of sulphur dioxide and sulphuric acid. A side product is sulphuric acid (93.9 wt%, Iiquid 30°C) which contains a small trace of sulphur dioxide. The mass flow of the sulphur trioxide productstream is 1106.6 kg/h and the mass flow of the sulphuric acid-productstream 286.0 kg/ho
3.1.3. Utilities
3.1.4. List of chemicals
Table 3.1.2 contains the chemicais, used in the sulphur oxidation process. Table 3.1.2: List of chemicals.[22],[231.
Nitrogen Oxygen 8ulphur- 8ulphur- 8ulphuric- 8ulphur dioxide trioxide acid
198 wt°/~ N2 O2 802 803
H
2804 8 boiling -196.0 -180.0 -10.0 44.8 279.6 444.6 point (OC) melting-
-
-175.0 16.9 10.4 110.2 point (Oe) densit~ 1.1 1.3 1460.0 1900.0 1843.0 1787.0 (kq/m ') molemass 28.0 32.0 64.1 80.1 98.1 32.1 (kg/kmole) MAC-
- 5-10 1 1 -(ppm) lethal dose-
-
3000 30 135 175.0 (ppm) prize 0.0 0.0-
- 130 165.0 (Dfl/ton) 153.2.
Precess descriptien
3.2.1. General process structure
The process of sulphur oxidation to sulphur trioxide is given in figure 3.2.1
1 I
11 111
2
Figure 3.2.1: Process structure of the S03 production Unit I: Drying of air with sulphuric acid.
Unit 11: Combustion of sulphur to sulphur dioxide.
Unit
111:
Oxidation of sulphur dioxide to sulphur trioxide.Unit IV: Condensation of produced sulphur trioxide.
IV
3
V 4
5
Unit V: Production of sulphuric acid out of waste sulphur trioxide and air clean-up Stream 1: Airfeed (4851.8 kg/h)
Stream 2: Sulphur feed (535.0 kg/h) Stream 3: S03 product (1109.5 kg /h) Stream 4: Off-gas (3991.3 kg/h)
Stream 5: Sulphuric acid 94% product (286.0 kg/h)
The different sections of the process will be discussed in 3.2.2.
3.2.2. Description of sections
-Air DryerThe air used for the oxidation of sulphur and sulphur dioxide is dried before entering the process because any remaining water will react with the sulphur trioxide to
sulphuric acid. The water is removed by absorption in sulphuric acid in an absorption column. The water that is removed from the air will be used later to produce
sulphuric acid out of waste sulphur trioxide.
The average relative humidity of air is 70%. The absorption column is although designed to remove all water out air with a relative humidity of 100%. This means that the air flow entering the absorption-column can hold a maximum of 96.4 kg/h water but will usually contain 66.8kg/h. The humidity is removed out of the air by a sulphuric acid stream (99.0 wt%, liquid, 30°C) of 1229.0 kg/ho The sulphuric acid
-Sulphur combustion
The sulphur is burned in a LURO® Rotary Sulphur Burner. The sulphur feed enters the burner with a flow of 535.0 kg/h (6 bar and 140°C). It is combusted with air which enters the bumer with a flow of 4785.2 kg/h (6 bar and 180°C). The sulphur is assumed to react with a 100% conversion to sulphur dioxide. The product gasses leave the burner at a temperature of 1057.9 °C (5.85 bar).
-Sulphur dioxide oxidation
The produced sulphur dioxide enters, after been cooled to 440°C, a fixed bed reactor where it will react with oxygen to sulphur trioxide. The reactor is made out of four fixed beds with intermediate cooling of reaction-gasses. The beds are made of a 10% vanadium pentaoxide catalyst mounted on a silica containing alumina support. Because sulphur dioxide oxidation is an exotherm ic process, the temperature of the gas increases. The equilibrium of the reaction moves from sulphur trioxide to sulphur dioxide and oxygen with increasing temperature and decreasing pressure. Therefore intermediate cooling is used to allow higher conversions. The water remaining from the air dryer reacts with sulphur trioxide in the first bed to sulphuric acid. This creates no serious problems as long as the temperatures are high enough to avoid
condensation.
The feed is 5320.2 kg/h (5.75 bar, 440°C). The gas-composition entering the reactor is: 20.1 wt% sulphur dioxide, 10.9 wt% oxygen and 69.0 wt% nitrogen. The
conversions and the temperatures in each bed are given in table 3.2.1
T bi 3 2 1
a
e
. . .
.
F Ixe d b d t e empera ures an conversIons t dbed number Temperature in Temperature out Conversion
(OC) (0C) (%)
1 440.0 647.8 72.0
2 440.0 507.3 95.0
3 425.0 435.4 98.5
4 400.0 402.7 99.4
-Sulphur trioxide condensation
The produced sulphur trioxide must be separated from the remaining air and sulphur dioxide. This is done by cooling the mixture (30°C) and increasing the pressure (25 bar). Because of possible condensation in the compressors, the condensation takes place in two steps, first at 30°C and 11.2 bar and second at 30°C and 25 bar.
Although the boiling point of sulphur trioxide is 44.8 °C, a quite large amount of sulphur trioxide would remain in the gas if the pressure is not increased, because of the low concentrations of sulphur trioxide in air. The temperature of the condensers
(30°C) allows for cooling with water and there is, at this temperature, no risk for solidification of the sulphur trioxide.
The feed of the condenser is 180°C and 11.2 bar. The feedgas composition is 24.9 wt% sulphur trioxide, 0.1 wt% sulphur dioxide, 0.0 wt% sulphuric acid, 5.9 wt% oxygen and 69.0 wt% nitrogen. Of the 1327 kg/h sulphur trioxide that enters the condensers, 840.9 kg/h condenses in the first condenser and 265.7 kg/h condenses in the second condenser. This means that 1106.6 kg/h is sent to the FSA-plant and 219.6 kg/h is sent to the sulphuric acid absorption column. In the two condensers, besides the sulphur trioxide, also 1.9 kg/h sulphur dioxide and 1.0 kg/h sulphuric acid condenses.
-Production of Sulphuric acid and gas clean-up
The waste-gases of the sulphur trioxide process are cleaned up before they are exhausted into the air. The remaining sulphur trioxide is absorbed in sulphuric acid (93.9 wt%) and reacts there with the water to sulphuric acid (99.0 wt%). The
sulphuric acid, which is used for the absorption, is drawn from the airdryer. The produced sulphuric acid is recycled to the airdryer.
Before entering the absorption column, the gasses out of the condenser are
expanded to 1.4 bar. The gas-composition of the entering gasses (4210.7 kg/h) is 5.2 wt% sulphur trioxide, 0.1 wt% sulphur dioxide, 7.5 wt% oxygen and 87.2 wt% nitrogen. The sulphur trioxide in the gas that leaves the column is neglectable. -Production of high-pressured steam
The heat produced in the reactors is recovered by the production of steam. This steam is at 40 bar and 410°C. Assumed is that there is steam-condensate available at 90
oe
and 40 bar. The condensate enters a steam vessel where the condensate is heated to 250oe.
The needed heat is supplied by condensing saturated steam. The vessel contains a mixture of steam and condensate at 250oe.
The heat-exchangers between and after the reactors draw condensate from the vessel and reboil this to saturated steam (250oe,
40 bar, vapour). One heat-exchanger superheats the produced saturated steam from 250oe
to 410°C. This superheating heat exchanger is placed directly behind the sulphur bumer and draws saturated steam from the steam vessel instead of condensate. The produced steam is used in theHF-destillation and any remaining steam is supplied to the steam-net present at the site.
3.2.3
Mass balance
In figure 3.2.2, the overall mass input and output is shown. S I hurtri oxi de up u,lphur feed Offgas Process S Air feed Sulph uricacid Figure 3.2.2: Overall mass input and output.
T bi 322 I
a
e.
.
: n- an d OUlpU s reams tt t
0 f h t e process. flow N2O
2 802 803 H20 H2804 S (kWh) (kglh) (kglh) (kglh) (kq/h) (kWh) (kglh) 8ulphur--
-
-
-
-
-
535.0 feed Air feed 3672.6 1112.4-
-
66.8-
-803-
-
1.9 1106.6-
1.0 -product Oft-gas 3672.6 314.0 0.6 --
-
-H2S04-
--
- 17.3 268.7 -product to waste treat--
-
3.9-
35640 1.1 -ment3.2.4 Simulating in ChemCAO
eThe whole process is simulated in ChemCAD with the SRK-model. However, ChemCAD is not capable of calculating the heats of reaction correctly which has a serious impact on the equilibrium. Therefore all the heats of reaction in each reactor have been calculated manually too. The condensers are designed manually because no adequate parameters could be found for the sulphur trioxide/air system. The absorption columns are also designed manually since ChemCAD is not sufticiently capable of calculating columns with sulphuric acid.
3.3 Equipment design and description
3.3.1 Air dryer
The humid air entering the air dryer is dried by absorption of the water into sulphuric acid. The sulphuric acid is diluted from 99.0 wt% to 93.9 wt%. The gas dryer is designed using design rules as described in Coulson & Richardson J1]. ChemCAO® was not capable of properly designing the column since ChemCAO posseses no suitable model for this equilibrium.
According to Comell's method, two values are estimated. If the factors (m:GmfLm)
and NOG are estimated and the ratio of the in- and outgoing concentrations is specified, the column parameters are easily calculated.
First the type and size of packing is chosen. For this tower, 38 mm ceramic Intalox saddle are thought to be the best cho.ice. From table 11.2 [1] Fp is read. Oensities
and viscosities of both vapour and liquid are taken from chemical data literature such as Perry [14]. A pressure drop of 20 mm H20/ m packing is chosen and the
1<.;-constant is read out of fig 11.44 [1] . The percentage flooding is calculated using the K;-constant. With equation 11.118 [1] the vapour mass flow per unit area is
calculated. The column area and diameter are calculated by dividing the mass-flow by the mass-flow per unit area.
To calculate the height of an overall phase transfer unit HOG, first the vapour and liquid Schmidt numbers are calculated. The Schmidt number is defined as:
Sc
:
=~
p·n
with Sc=
Schmidt number (-)~
=
viscosity (Pa.s) p=
density (kglm3)o
=
diffusivity (m2/s)(3.3.1)
The relationship between
HOG
and the individual film transfer units HL andHG
is given by:mG
mH OG := H G -+- • HL
Lm
with
HOG
=
height of the overall phase transfer unit (m)HG
=
height of the gas film transfer unit (m)HL
=
height of the liquid film transfer unit (m) m=
Slope of the equilibrium lineGm
=
molar flow rate of gas per unit area (mol.m·2.s·1)Lm
=
molar flow rate of liquid per unit area (mol.m·2.s·1)(3.3.2)
The heights
HG
andHL
are calculated using Cornell's equation 11 .110 and 11 .111[1]. The height of the column is calculated by multiplying the HOG by the number of steps needed,
NOG.
lurgi Sulphur Fumace secondary air rotary sulphur furnace -1--+- - - _. -primary air
Figure 3.3
.
1
lURO Rotary Sulphur Burner
belt transmission
Figure 3.3.2
burner shaft
electric
motor seal air
~
primary air ...--_ _ _ burner housing atomizer hood atomizing cup sulphur distributor3.3.2 Sulphur combustion reactor
In the first reactor, the sulphur is bumed with air to sulphur dioxide. In theory sulphur dioxide concentrations of about 20.5 vol% can be reached. But in practice excess air is supplied and the gasses that leave the reactor contain around 10-10.5 vol%
sulphur dioxide. This is done for two reasons:
1. to ensure complete combustion of the sulphur because the problems of condensing unbumed sulphur elsewhere in the plant are not worth risking. 2. by diluting the gasses the combustion temperature decreases. Too high
temperatures in the reactor may cause unacceptable NOx emissions.
Sulphur combustion in industrial oxygen or in oxygen-enriched air is, as a general rule, not economic on account of the high cost of oxygen, and it is not currently used on an industrial scale.
In order to combust sulphur with the air at nowadays standards, it is essential for the sulphur to be finely and evenly dispersed and intimately mixed with combustion air during buming. Today it is almost universal practice to use liquid sulphur at 140
oe
at which temperature the viscosity is low enough to enable it to be sprayed through nozzles.The reactor is essentially a fumace combined with a high-pressure-nozzle bumer. The fumace size and shape are very dependent on the type and efficiency of the bumer used. A picture of the fumace is given in figure 3.3.1. The fumace comprises a cylindrical steel shell lined with severallayers of insulating and refractory bricks. There are various ways in which liquid sulphur can be atomised to produce the fine spray needed for the most efficient possible combustion. The bumer used in the design is a rotary bumer in which the sulphur is distributed centrifugally from the edge of a spinning cup into a stream of air. The design of this Lurgi-developed LURO® Bumer is shown schematically in figure 3.3.2.
The sulphur passes into an open cup rotating at 4000-6000 rpm. The sulphur film on the intemal wall is dispersed form the edge by centrifugal forces and is further
atomised by entrainment in the so-called "primary air", which emerges at high
velocity through the narrow annular space between the rotating cup and the atomizer hood. "Secondary air', constituting of the majority of the combustion air, enters the fumace through an annular space between the atomizer hood and the burner casing. The system of adjustable air guide vanes in this space imparts a tangential
component to the flow of secondary air as it enters the conical combustion space between the burner and the main fumace chamber. The resulting spiral path of the gases through the fumace not only effectively keeps the flame central but at the same time prevents unvaporized sulphur droplets from impinging against the furnace wal!. A highly advantageous characteristic of rotary bumers is that the throughput can bevaried continuously from the rated capacity down to 20% capacity or less without any appreciabie change in atomising intensity.
The in- and outgoing streams of the sulphur burner are specified in Appendix C. The specifications of the reactor are given in Appendix E.
3.3.3 Sulphur dioxide converter
The sulphur dioxide from the sulphur-burner is converted to sulphur trioxide in a fixed bed reactor.
(3.3.3)
This reaction is an exothermic reaction and a lot of heat is produced which heats the gas-mixture. But with increasing temperature the equilibrium moves to the left. Therefore in the fixed bed, the temperature rises and the equilibrium is reached at a lower conversion. The equilibrium equation is:
with Kp = equilibrium constant (atm-o·s)
P(S03) = partial pressure of sulphur trioxide (atm) P(S02) = partial pressure of sulphur dioxide (atm) P(02) = partial pressure of oxygen (atm)
The Kp can be calculated using the Bodenstein and Pohl equation: 5186.5
IogK p
=
T +O.611·IogT-6.75 with Kp=
equilibrium constant (atm"c.s)T = temperature (K)
(3.3.4)
(3.3.5)
The problem can be solved by installing more beds with intermediate cooling. By installing four beds in the reactor, the conversion can come at the desired levels. High conversions are desired out of an environmental point of view instead of an economical. Modern installations are not allowed to emit huge amounts of sulphur dioxide. To avoid technical problems in the cleaning of the oft-gas, the reactor is designed to produce a high conversion.
The equilibrium of each bed can be calculated using formula (3.3.4) and (3.3.5).
Since the rate of reaction decrease substantially near the equilibrium, the conversion in each bed is chosen at areasonabie distance from the equilibrium. Severe
problems arose while calculating the reactor beds. Therefore the amount of catalyst for each bed is estimated using reliable literature data.
The catalyst beds are made of a 10% vanadium pentaoxide catalyst mounted on a silica containing alumina support. Oxidation on a vanadium catalyst is a
homogeneous reaction in a liquid melt of active components on both the internal and the extern al surfaces of an inert solid catalyst base. The reaction mechanism and the chemical structure of the active components have not been clearly defined.
According to the model of Mars and Maessen, the reaction takes place by the way of the intermediate steps shown in equation (3.3.6) and (3.3.7):
(3.3.6) (3.3.7)
The catalyst is irreversibly damaged above about 650°C. Below 400°C no stabie operation is possible due to the fact that below 400°C the active layer is no longer sufticiently melted. Therefore the temperature range in the reactor is between 400 and 650°C.
The vanadium catalyst is very insensitive to catalyst poisons. The ave rage service life is about 10 years according to catalyst manufacturers. Generally, the service life is determined not 50 much by progressive 1055 of activity as by catalyst losses incurred when filling and emptying the reactor and during routine screening. The in- and outgoing streams of the reactor are specified in Appendix C and the specifications of the reactor are given in Appendix E.
3.3.4 Sulphur trioxide absorption column
The remaining sulphur trioxide, in the oft-gas stream leaving the condensers and the turbines, is absorbed in sulphuric acid. The sulphuric acid is drawn from the air dryer. The sulphur trioxide in the oft-gas stream reacts with the water in the sulphuric acid (93.9 wt%) to more sulphuric acid. Therefore the concentration of acid leaving the column is increased to 99 wt%.
The column could not be properly designed in ChemCAO® since ChemCAO® possesses no suitable model for this equilibrium. Therefore the column has been design with the method in Coulson and Richardson [1] as described in paragraph 3.3.1.
The heat of reaction produced in the column is removed with a heat exchanger placed below the column. The heat exchanger uses water as a cooling medium. The in- and outgoing streams of the absorption column are specified in Appendix C and the specifications of the column and the heat exchanger are given in Appendix
E.
3.3.5
Sulphur dioxide absorption column
The sulphur dioxide which remains in the oft-gas stream leaving the sulphur trioxide
absor~tion column, is absorbed in water. The emission-limit of sulphur dioxide is 200 mg/m therefore 86% of the remaining sulphur dioxide must be absorbed. The
column has been design with the method in Coulson and Richardson [1] as described in paragraph 3.3.1.
The in- and outgoing streams of the absorption column are specified in Appendix C and the specifications of the column are given in Appendix E.
3.3.6 Nomenclature
DG
m HG HL HOGKp
Lm
m
NOG
P(02) P(S02) P(S03) ScT
p !lDiffusivity of vapour or liquid (m2/s)
Molar flow rate of gas per unit area (mol.m·2.s·1)
Height of a gas film transfer unit (m) Height of a liquid film transfer unit (m)
Height of overall gas phase transfer unit (m) Equilibrium constant (atm-o·s)
Molar flow rate of liquid per unit area (mol.m·2.s·1)
Slope of the equilibrium line
Number 6f overall gas phase transfer units Partial pressure of oxygen (atm)
Partial pressure of sulphur dioxide (atm) Partial pressure of sulphur trioxide (atm) Schmidt number
T emperature (K)
Density of vapour or liquid (kg/m3)
Viscosity of vapour of liquid (Pa.s)
3.4 Literature
[1] J.M. eoulson and J.F. Richardson. 'Chemical Engineering, volume 6, ~d edition',
Great Britain 1993
[2] U.H.F. Sander, H. Fischer, U. Rothe and R Kola, 'Sulphur, Sulphur Dioxide and
Sulphuric Acid', The British Sulphur Corporation Ltd., 1984
[3] MetalIgeselIschaft A.G., 'Werkwijze voor de katalytische omzetting in twee
contacttrappen van S02 in S03 enlof zwavelzuur', Nederlandse Octrooi 147105
1965
[4] Boc Limited, 'Katalytische oxydatie van zwaveldioxide', Terinzagelegging 7803639,1978
[5] J.M. Coulson and J.F. Richardson. 'Chemica I Engineering, volume 3, ~ ed.',
Great Britain 1994
[6] A. Simecek, J. Michalek, B. Kadlec and J.Vosolsobe, 'Löslichkeit von
Schwefeldioxid in Schwefe/säure', Chemie Ing. Techn. 1969
[7] Engelhard vanadium catalyst V-0701 T 1/8" data sheet
[8] Chemical Construction Company, 'Process for the production of sulphuric trioxide
and sulfuric acid', Patent 3455652, 1969
[9] Canadian Industries Ltd., 'Verfahren zur Herstellung von Schwefe/säure', Offenlegungsschrift 2255206, 1972
[10] J. Cathala, 'Process for the manufacture of sulfur trioxide', Patent 2510684, 1950
[11] EI. du Pont de Nemours & Company, 'Method and apparatus for the
manufacure of sulphuric acid', Patent 2075075, 1937
[12] EI. du Pont de Nemours & Company, 'Sulphur trioxide and Oleum, stro rage and
handling', USA
[13] RC. Reid, J. M. Prauznitz and B.E. Poling, 'The properties of gasses and
liquids, 4th ed.', U.S.A 1987
[14] R.H. Perry and D. Green, 'Perry's Chemica! Engineers' Handbook, fJh ed.',
Singapore 1984
[15] D.R Lide, 'CRC Handbook of Chemistry and Physics, 7fJh ed.', USA 1995 [16] J.M. Douglas, 'Conceptua! Design of Chemica! Processes', Singapore1988 [17] L.B.P.M. Janssen and M.M.C.G. Warmoeskerken, 'Transport Phenomena Data
Companion', Delft 1991
[18] K.R Westerterp, W.P.M. van Swaaij and A.A.C.M. Beenackers, 'Chemcia!
reactor design and operation', Amsterdam 1993
[19] S.M. Wales, 'Chemical reaction engineering handbook of so!ved problems', Amsterdam 1995
[20] J.J. Carberry and A. Varma, 'Chemica! reaction and reactor engineering', New Vork 1987
[21] J. Grievink, C.P. Luteijn and M.EA.M. Thijs-Krijnen, 'Handleiding
fabrieksvoorontwerp', TU Delft, 1994.
[22] Chemiekaarten
[23] RJ. Lewis sr, 'SAX's, dangerous properties of industrial materia!s', 8th ed., van Nostrand Reinhold, 1992.
4. Part three: Conversion to Fluosulphuric Acid
4.1. Basis of Design
In the previous chapters the production of sulphur trioxide and hydrogen fluoride is discussed. Now a closer look is taken at the part of the plant where these chemicals react to the final product: fluosulphuric acid. In two ways this subject is easier to discuss than the former subjects. In the first place, the pure reactants are now at our disposal; in the second place, the process is very weil described in a patent, found in literature [1] and in the basic data report of the Ou Pont fluosulphuric acid plant in La Porte [2]. This enabled us to examine this process in close detail.
In short, the process consists of a recycle stream of fluosulphuric acid, to which stoichiometric amounts of respectively sulphur trioxide and hydrogenfluoride are added. The following reaction takes place:
(4.1.1)
After cooling, a product stream is split from the recycle stream.
4.1.1. Feed
Two streams are fed to the fluosulphuric acid production. The first feed stream consist of 99.8 wt% pure sulphur trioxide. The mass flow rate of 803 is 1106.6 kg/ho The impurities are sulphur dioxode (1.9 kg/h) and sulphuric acid (1.0 kg/h). The S03 is fed from a tank storage that contains liquid sulphur trioxide at a temperature of 30
°C and a pressure of 1 bar. In the feed line to the F8A production, the liquid 803 is pumped to a pressure of 2 bar.
The second feed stream consist of 99.8 wt% liquid hydrogenfluoride at a temperature of 10 °C and a pressure of 1 bar. The mass flow rate of hydrogen fluoride is 275.4 kg/ho The only impurity is water (0.6 kg/h).
4.1.2. Product specifications
The product stream is liquid fluosulphuric acid of specified purity at a temperature of
38 °C and 1.85 bar. The composition of the stream on a mass flow rate basis is:
T bi 411 S
a
e
. . .
;pecl Ica lons T f 0 f th fl e UOSUlpl unc aCI I h . 'd pro uc S reamd t t .Name Formula Rate Concentration Specifications
(kglh) (wt.%) (wt.%) fluosulphuric acid HS03F 1377.6 99.4 98.5 sulphur dioxide S02 1.9 0.14 0.15 sulphur trioxide S03 1.7 0.12 0.30 sulphuric acid H2804 4.3 0.31 1.20 26
So the purity of the product is 99.4 wt%. The product contains no measurable amounts of water and hydrogen fluoride. As the product is strongly corrosive to metals, it will contain a small amount of iron (Fe) from the pipes and reactor.
All impurity concentrations are below the specifications set by Ou Pont. The product is stored in a double jacket storage tank, which is already present at the site.
4.1.3.
Utilities
The utilities available on site are the same as those summarised in chapter 2.1.3. Interested people may there satisfy their curiosity.
4.1.4.
List of chemicals
The following table contains the chemicals used in the production of fluosulphuric acid.
Table 4.1.2: List of chemicals [3],[4].
Hydrogen fluoride Sulphur trioxide Fluosulphuric acid (99.5 wt%) HF S03 HS03F boiling point 19.5 44.8 162.7 (OC) melting point -83.0 16.9 -89.0 .(OC) densit~ 958.0 1900.0 1725.0 (kQ/m ') molecular weight 20.0 80.1 100.1 (kQ/kmole) MAC 2.5 1 mg/m" 1 mg/m;j (ppm) lethal dose 774.0 30 mg/m,J 347 ppm (ppm) prize 1600 n.a. 1650.0 (Ofl/ton)
All three chemicals react violently with water. Hydrogen fluoride and fluosulphuric acid are strong acids and very corrosive, especially in contact with water. The possible dangers of these chemicals are considered very large; a safety analysis of the process is present in chapter 5.
4.1.5.
Literature
[1] U.S. Pat. No. 3,957,959. May 18,1976.
[2] Basic Data Report for Design of Fluorosulphuric Acid Plant. H.R. Williams.
E.1. Ou Pont De Nemours, La Porte, Tx, 1986.
4.2. Process description
4.2.1. General process structure
The process of reacting HF and 803 to fluosulphuric acid is given in the following figure. The flows are given in kg/ho
S03 feed S03 S 0 3 1106.6 Mixer S02 1.9 H ~O" 1.0 S03 S03 1176.3 rich S02 82.1 FSA FSA 57495.0 H~O" 181.1 HF feed HF
-+1
Mixer HF 275.4 Hp 0.6 FSA Reactor FSA recvc1e S03 69.7 S02 80.2 FSA 57495.0 H~O. 180.1I
FSA product Is03 71.4 SOa 82.1 FSA 58872.6 H~O" 184.4Figure 4.2.1: Process structure of H803F production.
The different sections of the process will be discussed below.
4.2.2. Description of sections
-503 mixing section Product Splitter FSA product S03 l.ï SOa 1.9 FSA 1377.6 H~O" 43In this section a controlled and measured liquid stream of 803 at 30 0 C and 2 bar contacts the recirculating fluosulphuric acid, which is at a temperature of 38 0 C and 2 bar. Since the temperature in this mixer is low, no Teflon coating is required. The material of construction is high grade carbon steel [1]. Backward flow in the sulphur trioxide feed is to be avoided. This is avoided by placing a pump in the sulphur trioxide feed line. The stream exiting the mixer is at 38.8 0 C and approximately 2 bar. The concentration of sulphur trioxide in this stream is 1.9 wt%.
- HF mixing / reaction zone
Immediately after adding hydrogenfluoride to the sulphur trioxide rich stream, reaction starts. The reacting stream is stirred by in-line mixers, to ensure good contact between the reactants. According to Ou Pont's basic data report [1], mixers and piping in the reaction zone should be constructed of 316 8tainless Steel. The length of piping in the reaction zone will be determined by the kinetics of the reaction. Thermodynamically, fluosulphuric acid is much more stabie than its constituents sulphur trioxide and hydrogen fluoride. Because there is a small stoichiometrie excess of sulphur trioxide (a mass feed ratio of 4.02 instead of a theoretical 4), the reaction will not stop until all hydrogen fluoride is converted.
The reaction is very exothermic. The value of the heat of reaction given by Ou Pont [1] is checked by thermodynamic calculation and is found to be correct. The
calculated value is 66.41 kj/mol. The value given by Ou Pont is 65.94 kj/mol. The first value is used.
The stream that exits the reaction zone is at a temperature of 52.8 °C and, as we assumed a sm all pressure drop over the reactor, a pressure of 1.95 bar. After the reactor an air cooler is placed, which cools the fluosulphuric stream from 52.8 °C to
38°C. The pressure drop over the air cooler is calculated by a method discussed by Brown [2]. The pressure drop is 0.1 bar. Air is used as cooling medium instead of water, because fluosulphuric acid becomes very corrosive in contact with water. If water is used as a cooling medium, a leak would result in a disaster.
- Recycle- and product stream splitter
The splitter controls the mass rate of the product stream. It will be made clear that a simple ratio control of this splitter is able to keep all streams in the process at a constant value.
4.2.3. Mass balance
From Figure 4.2.1 it is clear that two streams enter the process: a sulphur trioxide feed and a hydrogen fluoride feed stream. One stream of sufficiently pure
fluosulphuric acid leaves the process. Oue to the purity of both feed streams, the process does not need a separator and produces no waste stream of any kind. Table 4.2.1: In- and output streams of the process.
Stream FSA 803 802 HF H2804 H20 (kg/h) 803 feed 0 1106.6 1.9 0 1.0 0 HF feed 0 0 0 275.4 0 0.6 F8A 1377.6 1.7 1.9 0 4.3 0 product
4.2.4. Simulating in ChemCAOCC>
Simulating the reaction section in ChemCAD~ confronted us with several problems. ChemCADc does not have any data on fluosulphuric acid available in its data bank. Only after an elaborate literature study we were ab Ie to find the necessary data on fluosulphuric acid [3], [4], [5].
ChemCAOc is not able to calculate the correct heat of reaction for our liquid phase reaction, even if the correct heat of formation of fluosulphuric acid is available.
However, the chemical equilibrium is calculated correctly by ChemCAD~: the reactor simulated complete conversion of hydrogen fluoride to fluosulphuric acid.
So, as aresuit the mass balance for this part of the process is solved using
ChemCAD~ and the heat balance is solved manually.
4.2.5. Literature
[1] Basic Data Report for Design of Fluorosulfuric Acid Plant. H.R. Williams. E.!. Ou Pont De Nemours, La Porte, Tx, 1986.
[2] Brown R. (1978) Chem. Eng., NV 85 (March 27 th), 414. [3] Richards, G.W. and Woolf, A.A., J.Chem.Soc (A), 1967, 1118. [4] Chase et al., J. Phys. Chem. Ref. Data, Vol.3 , No. 2, 1974,422.
[5] Thompson, R.C. and Nickless, G., ed., Inorganic Sulphur Chemistry, Elsevier, Amsterdam, the Netherlands, 1968, 588.
