mek from n butene.pdf
TRANSCRIPT
F.V.O. Nr: 2693 Technische Universiteit Delft
Vakgroep Chemische Technologie
•
Verslag behorende
bij het fabrieksvoorontwerp
van
A.H. Amer
R.F. de Ruiter
onderwerp:
The production of methyl ethyl ketone
from n-butene
adres: Dr. H. Colijnlaan 187 A.M. de yonglaan 27 opdrachtdatum: 20-10-1986
2283 XG Rijswijk 3221 VA Hellevoetsluis verslagdatum: 12-07-1988
1
1.1 1.2 1.3 1.4 1.5 1.6 1.7
2
2.1 2.1. 1 2.1. 2 2.1. 3 2.1. 4 2.1. 5 2.1. 6 2.2 2.2.1 2.2.2 2.3 2.4 2.5 2.5.1 2.5.2 2.6 2.6.1 2.6.2
3
3.1 3.1.1 3.1. 2 3.1. 3 3.1. 4 3.1. 5 3.1. 6 3.2 3.3
4
5
6
Contents
Abstract
Conclusions and reco •• endations
General introduction
Uses and product ion Manufacture Choice of process Plant capacity Health and safety Feedstock Process description
Secondary butyl alcohol product ion
Butene absorber Liquification Absorption kinetics Material balance Heat balance and cooling Design Gas-liquid separator
Hydrolysis tank Material balance and design Heat balance
SBA stripper Caustic scrubber Sulfuric acid reconcentration unit
Reconcentration processes Drum design
SBA purification unit Liquid-liquid separator Azeotropic distillation unit
Methyl ethyl ketone product ion
Dehydrogenation reactor Convers ion of SBA Reaction thermodynamics Catalyst choice Kinetics of a Cu/Ni-catalyst Pressure influences Design
Hydrogen recovery MEK purification unit
Mass and heat balance, strea. data
Apparatus specifications
Cost esti.ation and econo.ics
References
page
1
2
3
3 4 5 6 6 6 7
9
9 9 9
10 10 12 13 14 14 15 16 20 21 21 22 23 23 24
27
27 27 28 29 30 33 34 35 35
39
52
67
73
.... __ ._-- _._----- --- - - --------
Abstract
In this preliminary design the production of methyl ethyl
ketone (MEK) from normal butene, with secondary butyl alcohol (SBA)
as intermediate, is described. This design is split into two parts.
In the first part SBA is obtained from n-butene by absorption in
sulfuric acid, followed by hydrolysis with water. Sulfurie acid and
SBA are separated in a stripper. The sulfurie acid is
reconcentrated and recycled to the absorber. The SBA is purified in
an azeotropic distillation unit, using diisobutylene as entrainer.
In the second part of the design, SBA is vaporized and fed to a
mul ti t ubular, isothermi c reactor, fi lIed wi th a Cu/Ni on S iO Zo
catalyst. The SBA is dehydrogenized, forming MEK and hydrogene The
hydrogen is purified and sold as a valuable by-product. The MEK is
purified in two fractionation columns and obtained with a purity of
99.1 wt"-%.
The capacity of the plant is 33,731 tons of MEK per year. An
economie evaluation shows that this plant can pay itself back
within approximately 1.5 to 2 years.
1
- - - - - ------------
Conclusions and reco •• endations
The extractive distillation unit, where SBA and water are
separated is simulated, using the UNIFAC group contribution method
for predicting activity coëfficiënts. This simulation can only be
used as an indication. To make an accurate prediction of the be
haviour of this unit, it is necessary to have reliable
thermodynamic data. The same problem occurs with the SBA stripper.
The influence of sulfuric acid on the equilibrium data could not be
forecasted and the assumptions made are rat her rigourous.
Although a compressor is attached, it is likely that n-butene
can be obtained in liquified state. The compressor covers 17% of
the equipment costs
equipment costs form
and because in the used economic model the
the base for obtaining the total capital
investment, this percentage has great effect on the economics of
the proces. Nevertheless a pay-out time of 1.5 years and an inter
Dal rate of return of 58.2% give a good indication for the expected
perspectives. This is due to the great difference between butene
costs and MEK selling prices. The price difference of f.200,-/t
between SBA and MEK can not justify the design of an SBA convers ion
plant only.
2
1 General introduction
1.1 Uses and production
Methyl ethyl ketone is one of the lowest priced solvents in its
boiling range and it is widely used as a solvent in a great variety
of coating systems. As a solvent for lacquers, MEK is particularly
advantageous because it provides low viscosity solutions at high
solid contents without affecting film properties. MEK is also used
as a dewaxing agent in the refining of lubricating oils and as a
solvent for adhesives, rubber, cement, printing inks and cleaning
solutions. It is used in vegetable-oil extract ion processes and in
azeotropic separation schemes in refineries [IJ. Furthermore it is
used in the pharmaceutical industry. Table(l-l)lists the main uses
of MEK for 1977 in the USA.
Table(l-l): Methyl ethyl ketone uses
Use
Vinyl coatings
Nitrocellulose coatings
Adhesives
Acrylic coatings
Miscellaneous coatings
Lube-oil dewaxing
Miscellaneous and export
Percentage
34
14
14
12
7
7
12
The output of MEK in the United States of America reached 27,000
tons per year in 1976 and the demand is expected to increase an
nually by 6 %. The situation is similar in Western Europe and in
Japan. The total annual production of MEK in Western Europe in 1976
was 220,000 tons. In Japan it was 65,100 tons.
The industrial importance of MEK is rising because the use of
solvents such as alkyl aromatics and branched ketones, which have
3
high biostability will become restricted for reasons of conserva
tion of the environment, and they can be replaced by MEK. In the
USA this is already alegal requirement [2J.
1.2 Manufacture
Methyl ethyl ketone can be manufactured by a direct oxidation of
n-butenes in aqueous solutions of palladium and cupric chlorides
[3 J :
+ ----)
It is also commercially available as a byproduct from liquid-phase
oxidation of butane to acetic acid.
In general MEK is produced by a two-step process from n-butenes.
The first step is the convers ion of n-butenes into secondary
butanol (SBA). In the second step the formed SBA is converted into
MEK, wether by oxidation or by dehydrogenation.
Secondary butanol can be produced by the hydration of l-butene
in the vapor phase by passage with steam over asolid catalyst
containing phosphoric acid and the oxides of metals as Zn, Mg and
Fe, at a temperature of 240°C and a pressure of 9.9 atm. [4], or
over a mixture of boric acid and phosphoric acid catalysts at 388°C
and 380 atm., with a maximum convers ion of 8.5 % per pass [5J:
+ -----)
About 10 percent of the reacted butene is lost by polymerisation.
Secondary butanol is usually produced by absorption of n-butenes
in sulfurie acid, followed by hydrolysis with water:
-----)
(-----
CH 3 -ÇH-C&H s + 3 H&O -----) OS03 H
CH 3 -ÇH-C&H s + 2 H&O OS03 H
The absorption of but ene can be carried out in 65 wt-% sulfurie
acid at 50-60 oC, in 75-80 wt-% acid at 30-50 oC and in 90-100 wt-%
acid at 15°C or below [4]. Gaseous butenes can be absorbed in 80
wt-% acid at a temperature of 43°C and atmospheric pressure [6J,
4
liquid butenes can be absorbed at a temperature of 38°C and a
pressure of 2-3 atm.(7].
The second step is dehydrogenation or oxidation of secondary
butanol to methyl ethyl ketone. The dehydrogenation of SBA can be
done in the liquid phase at a temperature of l50-250oC with
catalysts as raney nickel or copper chromite (8], and in the vapor
phase over copper or zinc catalysts at higher temperatures and low
pressures. The oxidation is done by air over copper or zinc oxides
at temperatures between 250 and 400°C.
Several other licenced methods for producing MEK are described
in literature (1]:
-Oxidation by acid dichromate,
peroxide or sodium perchlorate.
alkaline permanganate, hydrogen
-Free radical addition of acetaldehyde and ethylene:
free radical initiator -----------------------)
-Isomerization of butene oxide:
-Isomerization of isobutyraldehyde:
1.3 Choice of process
Most of the methyl ethyl ketone now being produced is obtained
from n-butenes in two stages: the sulfuric acid hydration of n
butenes to produce secondary butanol, followed by dehydrogenation
of the alcohol to ketone. Although sulfurie acid hydration is an
energy consuming process and corrosion aspects can not be underes-
timated, its technology has been proven for decennia and, when a
5
hydration plant is combined with a refinery or a naphta cracker
(what are also favorable combinations regarding the butene supply),
a major part of the required energy can be supplied from waste-heat
from flue gases. In the second stage the dehydrogenation is
preferabie to the oxidation, as the temperature regulation is
easier, the MEK yield is higher and hydrogen is formed as
byproduct.
1.4 Plant capacity
A design had to be made for a plant, capable to produce at least
30,000 ton MEK per year. To reach this target the feed of the plant
must be 23,347 tons per year of n-butenes (at a MEK yield of 100%).
The plant is designed to run continuous for 300 days per year (7~
hours per year). The actual butene feed is 26,457 t/yr and the
actual MEK production is 33,731 t/yr. The MEK is obtained with a
purity of 99.13 wt-% and the overall MEK yield from n-butene is
98.35%.
1.5 Health and safety
The toxic weight of methyl ethyl ketone in air is 200 ppm. For
the intermediate SBA this is 150 ppm. MEK is highly flammable
(flashpoint -lOC) and should be used with caution. The lower explo
sion limit is 1.8 vol-% in air and the upper explosion limit is 9.5
vol-% in air. For n-butene these limits are respectivily 1.6 and
9.7 vol-% in air and for SBA 1.7 and 9.8 vol-% in air. The electri
cal conductivity of MEK has a value of 2*10 7 pS/m, which means that
there is no danger for static charge build-up. Care should be taken
when MEK is stored for longer periods. Storage in carbon steel
tanks will lead to peroxide formation. Special alloys are available
which do not initiate this reaction.
1.6 Feedstock
Butylene
methylpropene
butene. The
These four
is the name of a mixture of four isomers: 2-
or isobutylene, l-butene, cis-2-butene and trans-2-
last three are referred to as normal- or n-butenes.
isomers and butane are treated as a C4 -group because
6
7 , j ,
they are of ten obtained as a mixture from cracked petroleum
fractions.
For the manufacture of secondary butyl alcohol (SBA) as inter
the product ion of methyl ethyl ketone (MEK) it is mediate for
necessary to have a feedstock in which the isobutylene is removed.
In electrophilic reactions isobutylene will react about thousand
times faster than the n-butenes and in our reaction scheme this
would lead to formation of tertiary butyl alcohol. However, this
difference in reactivity can also be used to separate the
isobutylene from the n-butenes. For this separation sulfuric acid
extraction can be used. Isobutylene can quantitativily be removed
in a solution of 45-60% HzSO. at 30°C.
Butane in the feedstock does not have affect on the but ene
absorption because it does not react with sulfuric acid. As in our
scheme unreacted butenes are recycled, inerts in the feedstock
would lead to accumulation and to prevent this, a part of the
recycle stream must be purged (e.g. to a furnace). r.; We assumed to have a gaseous feedstock at 1 atmosphere which
only containes n-butenes in their ~a~~~~l ~q~i]~b~~u~ distribution
at 300 K: 2 % l-butene, 9 % cis-2-butene and 89 % trans-2-butene
[24].
1.7 Process description
Gaseous butenes with a pressure of 1 atmosphere and a tempera
ture of 25°C are charged to a compressor, which is followed by a
cooler,
The
charged
to form
where liquification takes place at a pressure of 3 atm.
liquified butenes are mixed with 80 wt-% sulfuric acid and
to an absorption column. The acid reacts with the butenes
butyl sulfates and deprotonated secondary butyl alcohol.
The reaction is exothermic, and heat is withdrawn by cooling.
The conversion of butenes is practically complete (> 98 %).
Af ter the absorption stage the pressure is decreased to atmospheric
and residual butenes are removed from the product in a phase
separator and are recycled. The acid-sulfate mixture flows to a
hydrolyzer, where water is added and secondary butyl alcohol is
formed.
The hydrolyzate is fed to a column where the alcohol is stripped
from the diluted acid by means of life steam. Entrained acid is
7
captured in a demister and traces of acid in the alcohol-water
vapor are removed in a scrubber with diluted sodium hydroxide. The
scrubbed vapors are then condensed to form a crude containing water
and alcohol.
The diluted acid is reconcentrated in two stages and is recycled
to the absorption column.
The crude alcohol is, af ter separation in two liquid phases,
purified in a fractionation column. Diisobutylene (2,4,4-trimethyl
l-pentene) is added to the column as an entrainer to form a light
boiling ternary azeotrope in the top of the column, while alcohol
is withdrawn in the bottom. In a second column water is withdrawn
from the remaining mixture.
The secondary butyl alcohol is vaporized, preheated and charged
to a tubular reactor where dehydrogenation to MEK takes place. The
tubes are packed with a Cu/Ni on SiO z catalyst and are direct-fired
to maintain areaction temperature of 310°C. The reactor effluent
contains MEK, unconverted alcohol, hydrogen and a small amount of
water (the water comes with the alcohol from the fractionation
column). This effluent is condensed and charged to a phase
separator where the hydrogen is removed. The flue gasses of the
furnace are used for reconcentrating the diluted sulfuric acid.
The methyl ethyl ketone is purified in two fractionation
columns. In the top of the first column a mixture of MEK, alcohol
and a trace of water is withdrawn with a purity of MEK of 98.9
percent. The bottom product is charged to the second column. The
top product of the second column contains MEK with a purity of 99.3
percent and the bottom product contains the remaining alcohol which
is recycled to the reactor.
8
2 Secondary butyl alcohol product ion
2.1 Butene absorber
2.1. 1 Liquification
The liquification pressure of the mixture of butenes (89~ trans-
2-butene, 9~ cis-2-butene, 2% l-butene) is calculated by using the
Antoine equation for the vapor pressure:
(1)
where p is the pressure in mm Hg and T is the temperature in K and
A, Band Care to the vapor related constants. Values for these
constants are mentioned in appendix A-I . At a temperature of 25°C
the vapor pressure of the butene mixture becomes 1953 mm Hg (2.57
atm). The operating pressure in the column is fixed at 3 atm.
The gaseous mixture of n-butenes at atmospheric pressure and a
temperature of 25°C is compressed to 3 atm in a compressor and
liquified in a co~ The outlet temperature of the compressor is . .-XH.\' ( l' . 71°C, the actual ~ of the compressor 1S 73.72 kW. The condenser
duty is 1.76 MM kJ/hr (489 kW). These calculations have been done
with the program PROCESS on a mainframe computer and a printout of
the results is added in appendix A-2.
2.1. 2 Absorption kinetics
The relative rate of absorption of butenes into sulfuric acid
can be expressed by the following equation [9]:
x = l-exp(-K*t) (2)
where K is the absorption constant. K-values are mentioned for
gaseous and liquified butenes for various acid concentrations at
25°C [10]. For a sulfuric acid solution of 80 wt-% at a temperature
of 25°C, the absorption constant K has the value: K=33.48
xlO- 3 min- 1 for the above mentioned mixture of liquified butenes.
9
The relation between the convers ion percentage and the time is
shown in table (2-1):
Table (2-1): Conversion percentage of butenes at 25°C
in 80 wt-% sulfuric acid.
~ (min) 10 20 30 40 50 60 120
Conv. % 28.45 48.81 63.37 73.79 81.25 86.59 98.20
2.1.3 Material balance
180
99.76
For a conversion of at least 98% at 25°C, the residence time 0 which is needed is 2 hours. For equimolar amounts of sulfuric acid
~
and butenes it is necessary to have the following flow rates:
-Amount of butenes 3,742.6 kg/hr
-Density of liquid butenes at 25°C 602.09 kg/m 3
-Volume rate of liquid butenes 6.216 m3 /hr
-Amount of 80 wt-% sulfuric acid 8,032.54 kg/hr
-Density of sulfuric acid (80 wt-%) 1727.2 kg/m 3
-Volume rate of sulfuric acid 4.651 m3 /hr
2.1.4 Heat balance and cooling
During the absorption an excess of energy is released which has
to be removed as adequate as possible to prevent the temperature to
rise above 40°C. If the temperature of butene, in contact with 80
wt-% sulfuric acid, rizes above 60°C ,polymerisation will occur. To
prevent any polymerisation in the system the maximum reaction
temperature is set at 40°C.
It was not possible to determine the molar enthalpies for the
butylsulfate and the deprotonated SBA in the effluent of the ab
sorber and the assumption was made that they had the same value as
the molar enthalpy for normal SBA. During the absorption sulfuric
10
acid is diluted from 80 wt-% down to 54.6 wt-%. The involved heat
of mixing is calculated as if the acid is diluted with water. The
formed absorption products are to leave the column at a temperature
of 40°C. To achieve this temperature, it is necessary to withdraw
an amount of heat Q of 2166 kW. It is not possible to withdraw this
heat by the use of a jacket, filled with cooling water, because a
jacket can not provide anough area for heat transfer. To give an
idea for the required cooling area and the required amount of
cooling water, calculations were made for two different cases:
cocurrent and countercurrent flow of cooling water through pipes in
the column, made of stainless steel with a wallthickness d of 2 w mme
Foulingfactors: inside the pipes: hf(in) = 5.7 kW/mz.oC for
treated cooling water and outside the pipes: hf(out) = 2.8 kW/mz.oC
for inorganic liquids (12].
Heat conductivity coëffiënt for stainless steel:
W/m.oC.
The overall heat transfer coëfficiënt U becomes:
d + ---~-- +
À ss
U = 1538 W/m z . oe
(3)
À = 17 ss
If T(in) and T(out) are the temperatures of respectivily incom
ing and outgoing product streams and t(in) and t(out) are the
temperatures of respectivily incoming and outgoing cooling water
streams, the logarithmic mean temperature difference ~Tln follows
from:
(4)
for countercurrent cooling and:
= (T(in)-t(in»-(T(out)-t(out» -----î~-!I!~I=!I!~I==---------
T(out)-t(out)
(5)
for cocurrent cooling.
11
The required heat transfer area A can be obtained from:
Q A = ------ (6)
U . .1T ln
In table (2-2) .1T ln , cooling area A and required amount of
cooling water are mentioned as function of the outgoing cooling
water temperature.
table (2-2): .1T ln , cooling area A and required amount of
cooling water for co- and countercurrent cooling
water flow
t(out)
( Oe)
21
22
23
24
25
26
27
28
29
30
t(c.w.)
(m 3 /hr)
1861
931
620
465
372
310
266
233
207
186
countercurrent cocurrent
9.94 142 10.49 134
8.96 157 10.15 139
7.82 180 9.81 144
6.34 222 9.46 149
9.10 156
8.74 161
8.37 168
8.00 176
7.61 185
7.21 195
As can be seen from table (2-2) cocurrent coo1ing is preferabie
to countercurrent cooling. With increasing t(out) the required
cooling water flow t(c.w.) decreases while the required cooling
area increases.
2.1. 5 Design
With specific data about cooling water costs and heat transfer
area costs one can derive an optimal design. However, we assumed
12
- --------- - - - ---------------- --
that a 6Tln
of 8°C is the minimum acceptable driving force for
sufficiënt heat transfer and this fixes the cooling area at 176 mZ
and the cooling water flow at 233 m3 /hr. Another criterion is the
minimum allowable water velocity in the tubes. This velocity must
be above 0.7 mis to prevent fouling inside the tubes [40]. To
attain this velocity, the water must flow through a total, radial
tube surface of 233/3600/0.7 = 0.0925 mZ • Assuming a total of n
tubes, each with a height h, in the column, gives us the tube heat
exchange area A and the radial tube area A' as function of the tube
radius r:
A = 176 = 2 * n * r * h n
0.0925 A' = ------ = n * r Z n
(7)
(8)
The liquid butenes and the sulfuric acid are fed together in
the bottom of the column with a total volume rate of 10.867 m3 /hr.
With a residence time of 2 hours, the minimal required volume is
21.734 m3 • A column with a height of 13.7 mand a diameter of 1.5 m
provides a total volume of 24.210 m3• With the tube height h fixed
to 13.7 m, eq.(7) and eq.(8) can be solved and give us the number
of tubes n = 142 and the tube radius r = 0.0144 m. The total tube
volume Vtt becomes:
Substracting this value from the total column volume gives a
remaining absorber volume of 22.566 m3• This volume provides a
residence time for the butene-acid mixture of 2 hours and 4.6
minutes and a maximum butene absorption of 98.48% at 25°C.
At 40°C the absorption constant K is not known, but it can be l
assumed that absorption at that temperature will be complete.
2.1. 6 Gas-liquid separator
Af ter the absorption column the pressure is reduced to atmos
pheric and although but ene absorption is considered to be complete,
a gas-liquid separator is attached for removal of small amounts of
13
unreacted gases. We assumed these gases to be butenes and recycle
them to the entrance of the compressor. If the feedstock, however,
containes small amounts of inert ia as butane, a part of the recycle
is to be purged to prevent a build-up of these inert ia in the
absorber.
In general
1iquid. The
gravity is used for the separation of gas from
maximum horizontal vapor velocity U in the separator v is calcu1ated with the fo1lowing equation [21]:
o 5
U = 0.035 ( (Pl-p )/ P ) v v v (9)
where Pv and PI are the densities of respectively vapor and
liquid (kg/m 3 ). For our system the maximum vapor velocity becomes
0.53 mis. We want to remove a maximum of 2% of the initial amount
of butene, what results in a gas flow rate of 0.008 m3 /s. The
minimum
between
must be
diameter
area
the
20%
is
gas bubbles
the minimum
for vapor passage then becomes 0.015 mZ • The height h
top of the (horizontal) vessel and the liquid level
of the vessel radius R. Using this data, the vessel
calculated at 0.60 m. With a slip velocity for small
of 1 cm/s, the residence time becomes 54 seconds and
vessel volume for the liquid only 0.147 m3• Together
with the required gas volume, the total vessel volume becomes 0.164
m3 and the vessel length 0.60 m.
2.2 Hydrolysis tank
2.2.1 Material balance and design
Af ter the absorption of n-butenes in sulfuric acid the liquid
contains partially deprotonated SBA and secondary butyl sulfate.
Both components are completely and instantaneous converted into SBA
when excess water is added to the liquid. The sulfuric acid is
di1uted from 36.8% by moles (80 wt-%) down to 6.8% by moles (30 wt
%). At this dilution all intermediates are converted to SBA.
The feed of the hydrolysis tank contains 65.5 kmo1es/hr HzSO.,
65.5 kmoles/hr SBA and 46.8 kmoles/hr water. This represents a
total flowrate of 11,774.5 kg/hr. The density of this mixture is
derived with the following equation:
14
(10 )
Because we have to deal with highly corrosive sulfurie acid, a
hydrolysis tank is designed in which the fluid is not mixed by an
agitator with a shaft and inevitable seals, but in which the liquid
is mixed by the impuls of the incoming water stream. Racz et.al.
[13] stated that the mixing time of an aqueous solution in a tank
with approximately equal diameter D and height H can be calculated
with the following equation:
where:
D = tank diameter
d = nozzle diameter
v = velocity of the water in
t = m mixing time
With the following data:
-Density of productstream
-Flowrate of productstream
-Volume rate of the water to
dilute the acid to 30 wt-%
-Assumed nozzle diameter (2 inch)
-Assumed tank diameter
we obtain the following results
-Mixing time (t ) m
-Residence time (1.5*t ) m -Volume of the tank
-Height of the tank
2.2.2 Heat balance
(11)
(m)
(m)
the nozzle (mis)
(s)
1370 kg/m 3
8.542 m3 /hr
14.483 m3 /hr
0.0508 m
0.5 m
15.16 s
22.73 s
0.145 m3
0.740 m
Wh en sulfurie acid is diluted with water a large amount of
dilution heat is involved. It can roughly be estimated that in the
15
feed one mole of HZ S04 is solved in two moles of water. In the
product stream leaving the hydrolysis tank however, one mole of
HZ S04 is solved in thirteen moles of water. The molar enthalpy for
a mixture with an acid-water ratio of one to two is -204.55
kcal/mole
is -211,19
hydrolysis
HZ S04 and for an acid-water ratio of one to thirteen it
kcal/mole HZ S04 [19J. By diluting the acid in the
tank an excess of 6.73 kcal/mol HZ S04 (28.20 kJ/mol) is ~ -released. The total heat product ion becomes:
65.5 kmoles/hr HZ S04 * = =
1.847*106 kJ/hr
513.11 kW
The feed enters the hydrolysis tank with a maximum temperature of
40°C. If we assume the temperature of the water stream entering the
tank to be 25°C, the temperature of the productstream leaving the
hydrolysis tank is 51.4°C . Af ter dilution all butylsulfate and
deprotonated butylalcohol is converted into SBA and there is no
danger for polymerisation of the butene derivates. The product
stream can now be heated to 91°C (boiling temperature of the water
SBA azeotrope at 1 atm.) and fed to a stripper where SBA and acid
are separated.
2.3 SBA stripper
The product stream leaving the hydrolysis tank is a mixture with
86.34 mol-% water, 6.83 mol-% secondary butyl alcohol and 6.83 mol
% sulfuric acid. In this mixture acid and SBA have to be separated
from each other. It was not the intens ion to obtain one of the
components in its pure form. It was assumed that sulfuric acid, due
to its high boiling point (338°C) and due to the fact that it is
dissociated in water, did not take part in the vapor-liquid equi
libria of SBA and water. With this assumption only the binary
system SBA-water is left.
To define the number of equilibrium stages in the stripper, the
grafical method of McCabe-Thiele is used. The binary system is
described with the data in fig.(2-1) [25J. A part of this figure is
16
magnified and presented in fig.(2-2), together with the q-line, the work
seen line and the equilibrium stages which are obtained. As can be
in this figure, the azeotropic vapor separates in two liquid phases and
point (x sba
distillation can not
= 0.1 40 , Ysba = 0.396).
go beyond the first separation
(1) 2-BUTANOL
(2) WATER
+++++ ANTOINE CONSTANTS (1) 7.47429 1314.188 (2) 8.07131 1730.630
PRESSURE- 760.00 MM HG
CONSTANTS: A12
MARGULES 3.9182 VAN LAAR 3.7964 WILSON 11814.8851
NRTL 639.8173 UNIQUAC 350.171l7
EXPERIMENTAL DATA T DEG C Xl Yl
87.80 0.11110 11.36211 87.69 1l.1l2411 11.38211 87.911 11.31111 1l.39611 87.1111 0.3320 11.3960 87.IlII 11.3619 11.39611 87.19 11. 4781l 11.4999 87.29 11.51411 9. 4 lil 11 87.4Il 11.5629, 11.42211 87.59 11.58411 11.42611 87.611 0.61140 11.4360 87.70 11.6520 0.45011 88.10 0.6840 11.4640 88.10 0.71100 0.48411 911.20 0.860" 0.6219 92.70 0.91411 0.7160 93.80 0.93110 0.7580 95.80 11. 961111 0.8400
MEAN DEVIATION:
MAX. DEVIATION:
C4H 190
H20
REG ION +++++ 186.500 25- 120 C 233.426 1- 190 C
1.al3 BAR
A21 ALPHA12
1. 2808 1.4144
1643.6524 2491. U63 0.4385
309.5428
MARGULES VAN LAAR WILSON DIFF T
-7.32 -3.78
2.114 1. 95 1. 86 1.97 2. lil 2.27 2.32 2.34 2.08 2.10 1. 85
-1. 02 -1.36 -1. 21 -1.10
2.27
7.32
1.00
1 0.80
0.'0
YI
D.40
0.10
0.00
DIFF Y1 DIFF T
1l.1996 -3.33 11.9946 1.12
-11.11450 1. 56 -11.11369 1. 53 -0.9265 1.51
11.9961 1. 58 11.9199 1. 69 11.9296 1. 61 9.9397 1. 58 11.9367 1. 53 0.0352 1.15 0.0337 1.12 0.11442 0.87
-11.0045 -1. 30 -11.9249 -1. 25 -0.0239 -1. 91 -0.0277 -0.84
0.9417 1. 44
0.1906 3.33
.c~ lL V
/ V
V
DIFF Y1 DIFF T
11.11763 11.38 -1l.II257 9.22
0.0063 0.21 11.11112 0.23 0. U61 11.26 II.92U 11.35 9.9241 9.411 11.9237 11.47 11.112112 0.49 11.0224 0.511 Il.0134 0.31 Il.0983 9.44 0.1ll76 11.28
-0.1ll89 -0.80 -0.9258 -0.56 -0.9210 -0.33 -0.9203 -11.29
11.11219 0.38
9.1l764 C.81l
~ / I
lL 'f V " lL lL
~ K<
NRTl Y· -I - 51.95 Y· -. - 5.12
O~ O~ O~ O~ O~ I~
XI •
DIFF Y1
-".U22 9.0979
-0.0054 -0. ""82 -11.9123 -9.11289 -11.11269 -".9274 -9.9299 -11.11263 -0.93112 -11.11303 -11.9183 -0.0154 -0.11194 -0.11937 -11.111137
II.1ll74
II.1l303
DIFF T
-2.22 1. 79 11.56 9.57 11.58 9.62 9.64 9.66 11.66 0.64 0.39 9.49 11.31
-0.82 -11.51 -0.26 -0.211
0.70
2.22
figure (2-1) McCabe-Thiele diagram for the
system SBA-water at 1.013 bar
17
NRTL UNIOUAC DIFF Yl DIFF T DIFF Yl
11.9474 -3.53 9.9819 -0.9394 1. 96 -9.9243
11.11115 1. 64 II.0U1 11.0102 1. 61 1l.9957 9.91172 1. 59 9.9193
-9.9115 1. 66 9.U59 -9.11115 1. 79 1l.9195 -11. U511 1.72 II.92U -0.U911 1. 70 0.U7l -Il. U68 1.66 0.1l198 -0.0237 1.30 9.11121 -0.0258 1.28 9.91189 -11.0147 1.114 1l.1ll76 -0.0158 -1.14 -0.9157 -11.0091 -1.14 -0.0230 -11.0018 -Il.92 -1l.9186 -11.99118 -Cl.79 -1l.9188
0.0165 1. 50 0.9193
0.0474 3.53 0.9819
r Ysu
0.3
0./
1."/
..."
)tSBA
tI./O
figure (2-2): part of McCabe-Thiele diagram
from fig. (2-1)
The separation configuration is as follows:
over the top the binary azeotrope of SBA and water is withdrawn.
Practically all alcohol is withdrawn this way.
- the bottom product consists only of water (and acid).
- there is no reflux and no condenser in the top.
there is no reboiler. Vapor and energy are supplied by means of
steam injection in the bottom of the column.
The slope of the equilibrium line for
* "'sba * Psba K = 1 = -----------x p
x ~ 0 is given by: sba
(12 )
At 100°C, P:ba = 771.3 mm Hg, p = 760 mm Hg and "'sba = 51.95. The
K-value becomes 52.72. If we want to evaporate 65.5 kmol/hr SBA, an
energy of 758.4 kW is required. If steam of 1900C and 3 bar is
converted to water of 100°C and 1 bar, the enthalpy change is
42.577 kJ/mol. For SBA evaporation an amount of 64.12 kmol/hr steam
is to be condensed. To form an azeotrope with molefraction SBA = 0.396, an amount of 99.9 kmol/hr water vapor is required. A total
feed rate of 164 kmol/hr steam of 190°C and 3 bar is sufficiënt to
18
strip the SBA from the water-acid mixture. This implies a vapor
flow V in the stripper of 164 kmol/hr and a liquid flow L of 957.5
kmol/hr. For x sba factor S becomes:
S = K * ~ = 9.02
< 0.005 the K-value is constant and the strip
(13)
For constant S, the fraction f of not stripped SBA on a tray,
compared with N trays above this tray is calculated with:
(14)
The xf
= 0.0733 and as can be seen in fig.(2-2), af ter two
stages the x decreased to 0.004. In table (2-3) the compositions of
liquid and vapor are given for each tray. The trays are numbered
from the top down.
table (2-3): Tray number N and SBA fraction in liquid (x)
and vapor (y).
N x y
1 0.073 0.396
2 0.040 0.395
3 0.004 0.211
4 4.0e-4 0.021
5 4.4e-5 2.3e-3
6 4.8e-6 2.5e-4
7 5.4e-7 2.8e-5
8 6.0e-8 3.2e-6
The number of equilibrium stages is 8 and with an assumed (low)
Murphree tray efficiëncy of 60% the actual number of trays used in
the column is 13.
19 - ---- -
2.4 Caustic scrubber
If the demister on the top of the alcoholstripper fails, the
entrained acid-mist (max. 0.05 kgf kg vapor) must be removed by
another technique. This is necessary to prevent deactivation of the
catalyst used for the convers ion of SBA in MEK. This catalyst is,
like most catalysts, sensitive for small traces of sulfur in the
reactor input stream. The vapor is therefore scrubbed with a
diluted sodium hydroxide solution. The maximum acid-mist flow is
0.05*6653 kg/hr = 332.65 kg/hr. This mist contains maximal 28.55
wt-% acid (acid concentration in feed stripper), so a maximum of 97
kg/hr HZ S0 4 has to be removed. For this a NaOH-solution (9 wt-%)
flow of 465.3 kg/hr is needed. The diameter of this column, based
on 70 percent of the flooding velocity, is 1.0 m.
20
2.5 Sulfuric acid reconcentration unit
2.5.1 Reconcentration processes
Sulfuric acid acid reconcentration processes can be classified
in high-temperature processes, operating at atmospheric pressure
and in vacuum processes, operating at reduced temperatures [15].
High temperature processes have their major use in reconcentrating
acid with organic contaminants, which must be reduced to the lowest
possible level. For large scale concentration of relatively clean
acid the vacuum system is expected to be the process of choice,
because of the minimum air pollution possible. For reconcentrating
the sulfuric acid leaving the acid stripper and which contains a
small amount of secondary butanol, is choosen for the Chemico drum
concentrator as a high-temperature process [16J,
,
coo ..... -J. . , ....---__, :wd,f,f:i ,n
., .... IICI •
... ~.L. ""' ••
•• oovc, ac .. Hw .... , _,t •• eo.cI.' •• '''.'
Figure 2-3 Simplified flowsheet of Chemico drum concentration process.
21
The ehemico drum concentrator is used for concentrating sulfurie
acid solutions up to 93 wt-%. In this process, as shown in figure
(2-3), hot furnace gases are contacted with the acid in a serie of
vessels arranged countercurrently. The gases are blown onto the
liquid at approximately the liquid level through silicon iron dip
pipes and the vapors leaving the concentrator are scrubbed in a
venturi scrubber. The operating temperature is reported to be about
50°C below the atmospheric boiling temperature of the actual
mixture.
2.5.2 Drum design
It is necessary to use two drums to reconcentrate the sulfurie
acid coming from the acid stripper from 28.55 wt-% to 80 wt-%. In
the first and largest drum a reconcentration from 28.55 wt-% to 50
wt-% is achieved. In the second drum the remaining acid stream is
concentrated upto 80 wt-%.
First drum:
The reconcentration from 28.55 wt-% acid to 50 wt-%
-The boiling temperature for
50 wt-% acid solution
-Operating temperature
-Amount of water to be vaporized
-Heat required for evaporating water
-Heat of mixing (to be added)
-Tot al amount of heat required
(for the first step)
Second drum:
123
73
9,646.9
6.234
0.109
6.343
The reconcentration from 50 wt-% acid to 80 wt-%
-Boiling temperature for 80 wt-%
-Operating temperature
-Amount of water to be vaporized
-Heat required for evaporating water
-Heat of mixing
-The total amount of heat required
(for the second step)
22
196
146
4819.5
2.847
0.546
3.393
oe
oe
kg/hr
MW
MW
MW
oe
oe
kg/hr
MW
MW
MW
The total amount of heat required for reconcentrating the acid
stream is 9.736 MW.
2.6 SBA purification unit
2.6.1 Liquid-liquid separator
Wh en the SBA-water vapors from the caustic scrubber are con
densed, the formed liquid tends to separate into a light organic
phase and a heavy inorganic phase. The upper liquid layer has a
mole fraction x b 1 of 0.460 (77.8 wt-%) and the lower layer has s a,u a mole fraction xsba,ll of 0.040 (14.6 wt-%). This separation is
obtained in a liquid-liquid separator and occurs under the in
fluence of gravity, owing to the difference in density between the
two liquids [22J. Horizontal drums are generally used for this
separation. The required residence time t (min.) can be ap
proximated with the formula:
(15 )
with ~ the viscosity of the dispersed phase (cP) and PIl and Pul
the densities of lower and upper layer respectivily (g/cm 3 ). The
dispersed phase is the heavy, water-rich, phase and the viscosity
of water at 90 0 e is 0.3147 cP. At 90 0 e the densities of SBA and HzO
are respectivily 0.78347 g/cm 3 and 0.96534 g/cm 3 • the density of
the upper layer is calculated as:
x sba (wt-%) * Psba + x h 0 (wt-%) * Psba = _______________________ A ______________ _ 100 (16)
and has the value 0.8238 g/cm 3 • The lower layer density has the
value 0.9388 g/cm 3 • The required residence time is t = 8.21 min.
With a total flow rate of 1.832 kg/s, what is equal to 0.0022 m3 /s,
a minimum separator volume of 1.085 m3 is required. With a length
diameter ratio of 4, the separator diameter is fixed at 0.70 mand
the length at 2.80 m.
23
2.6.2 Azeotropic distillation unit
In figure (2-4) are two McCabe-Thiele diagrams presented, both
for the binary system HzO-SBA at 1.013 bar. One predicts a
heterogeneous azeotrope [25] and the other a homogeneous azeotrope
with liquid-liquid separation beside the azeotrope [26].
1.00
1 0.10
0.10
YI
0.40
o.ZO
0.00
r L 7
V /
V
V1 / I
/ f / ~
V / ~ ~
NRTL Y· -I - 51.95 Y· -. - 5.12
1.00
1 o.eo
0.80
YI
0.20
0.00
~
V
A V
/ V
VI / /
V V
/ y'
V A ~
V
NRTL Y· -I -
71.31 Y· -. - 5.05
0.00 0.20 0.40 0.10 0.10 1.00 0.00 o.ro 0.40 0.10 0.10 1.00
XI .. XI ..
figure (2-4): two different McCabe-Thiele diagrams for the
system SBA-water at 1.013 bar
In theory it is possible to separate SBA and water if they form
a heterogeneous azeotrope and not if they form a homogeneous
azeotrope. Furthermore the difference in boiling points is only
0.5°C and separation by normal distillation is for this reason only
very difficult. To make an SBA-water separation possible, an or-
ganic solvent (entrainer) can be added to the mixture, which forms
a light-boiling ternary azeotrope and is by this way able tobreak
the azeotrope. If the right amount of solvent is added, in one
fractionation column the mixture can be split in SBA and a mixture
with azeotropic composition, while in a second column the mixture
is split in water and again the azeotropic mixture. The tops of
both columns are connected with a decanter, where the condensed
azeotrope is splitted in a light organic layer and a heavy inor
ganic layer. Both layers are then recycled as reflux to the
columns. In table (2-4) four entrainers are mentioned with the
properties of the azeotrope they form with SBA and water. As can be
seen, diisobutylene (2,4,4-trimethyl-l-pentene, further referred to
as DiiB) forms an azeotrope with the smallest amount of water in
24
the organic layer and the smallest amount of SBA in the inorganic
layer.
ComponenlS Azeotrope:
~ . Percent composition Relative Spc:cific BP. . BP. volume of gravity
Compounds ·C ·C In azeo- Uppe:r Lower layers of layers .. trope: layer layer at 2o-C or azeotrope:
a. 2-Butanol 99.5 85.5 27.4 ~1.7 4.6 U 86.0 U 0.858 b. 2-Butyl acetate 1122 52.4 62.3 0.6 L 14.0 L 0.994 c. Water 100.0 20.2 6.0 · 94.8
a. 2-Butanol 99.5 86.6 56.1 65.0 10.0 U 86.0 U 0.816 b. Butyl ether
. . 1420 19.2 23.0 0.2 L 14.0 L 0.981
c. Water 100.0 24.7 12.0 89.8
a. 2-Butanol 99.5 67.0 · b. Cyc10hexane 81.0 c. Water 100.0
a. 2-Butanol 99.5 77.5 19.0 20.0 9.0±1 U 92.0 U 0.736 b. Diisobutylcnc 1026 70.0 78.8 0.5 L 8.0 L 0.987 c. Water 100.0 - ILO ; 1.2 91.0± I
table (2-4): ternary azeotropes, containing water and SBA
A computer program, provided by Magnussen et. al. [34], is used
to do the separation calculations. The algoritm of this program is
based on the separation calculations as presented by Naphtali and
Sandholm [35]: the equations of conservation of mass and energy and
of equilibrium are grouped by stage and then linearized. These
linearized equations are then solved simultaneously. Solution
convergence is obtained by the Newton-Raphson method. In the
program energy balances are not taken in account, but equimolar
overflow is assumed. The program uses UNIQUAC binary parameters to
predict activity coëfficiënts. These parameters were obtained with
the UNIFAC group contribution method. Program output for the
columns T23 and T29 plus the obtained UNIQUAC parameters are
presented in appendix A-4. The value for the molar heat of evapora-
tion of DiiB was not available and in the energy balance it is
given an arbitrary value Q.
25
In the figures (2-5) and (2-6) the component profiles in resp.
column T23 and column T29 are presented:
lIale fractian 1.8
8.5
DUB
SBA
H20 8.8*-~--~~~~~~~~~==~-,--~~~=-~-4
1 2 3 4 5 Ei 7 B 9 18 11 12 13 14 15 H. tray na.
figure (2-5): component profile for column T23
(stage 1 is in the bottom)
.ale fractlan 1.8
8.5
8.8l-~~~---+--~--~==~~~~-=~==~--~ 1 2 3 4 5 7 B 18 11 12
tray no.
figure (2-6): component profile for column T29
(stage 1 is in the bottom)
26
3 Methyl ethyl ketone product ion
3.1 Dehydrogenation reactor
3.1.1 Convers ion of SBA
There are basically two paths to convert SBA into MEK. One path
is partial oxidation with oxygen:
SBA + i Oz -----) MEK + HzO
This reaction is exothermic and a very good temperature control is
essential
CO, CO z ,
sufficiënt
oxidized
reaction
to prevent uncontrolled reactions in which byproducts as
butenes and other volatiles are formed. Even with a
temperature control, a large amount of the alcohol is
to HzO, CO and CO z . By using a catalyst as zinc-oxide the
temperature can be decreased to about 300·C and the yield
of MEK from SBA can be increased to 75-80 percent. However, a large
amount of the feed is turned into useless products which have also
to be separated from the MEK.
The second path is dehydrogenation of SBA by use of a catalyst:
SBA _E~!.!._) ( ______ MEK + Hz
This
the
reaction is endothermic and the maximum convers ion depends on
equilibrium constant of the reaction. Because energy has to be
added, the temperature control is much easier. Furthermore hydrogen
is formed as a valuable byproduct. This hydrogen is of a high
quality because it doesn't contain non-condensables.
Depending on the used catalyst, undesired byproducts can be
formed due to selfcondensation of MEK. These byproducts are of ten
unsaturated Ce-ketones like 3-methyl heptene-3-one-5, which are the
precursors of polymerisation and coking on the surface of the
catalyst, resulting in a rapid decreasing of the catalyst activity.
It is also difficult to separate these byproducts from the crude
MEK.
of
In this
the easy
design is choosen for a dehydrogenation of SBA because
temperature control, the formation of high quality
27
hydrogen as byproduct and because a catalyst was found that com
bined good activity and stability with a selectivity of 100% for
MEK.
3.1.2 Reaction thermodynamics
The dehydrogenation of SBA into MEK is a gasphase equilibrium
reaction:
with:
K ____ E ___ > SBA <________ MEK + Hz
p (ME K) * p ( Hz) K = p ---p(SBA)----- (17 )
Kolb and Burwell [17] derived three equations in which Kp' 6HTo
and
6STo
were found as function of the temperature (T in K):
log K -2790 + 1. 510 * log T +1.865 (18) = -----p T
6HTo
= 12770 + 3.0 * T (cal/mol) (19)
6STo
= 11. 54 + 6.908 * log T (cal/mol/K) (20)
In figure (3-1) the convers ion of SBA at equilibrium is plotted
as function of the temperature. Note that at a temperature of 200°C
the maximum convers ion is ~nly 60% and at 300°C the maximum conver
sion increases upto 93%. For a satisfying convers ion without a
large SBA-recycle stream, the reaction temperature must be above
300°C.
28
The
SM cOllYllra i on 1.8,---------------:=::::===;-8.9
8.8
8.7
8.6
8.5
8.4
8.3
8.2
8.1
8.8+-==~----_+----------+_--------__ --------_+ 8 188 4B8
figure (3-1): maximum feas ib Ie SBA convers ion ai , Q.tw.. as function of the temperature
3.1.3 Catalyst choice
gas phase dehydrogenation of SBA is supported by
heterogeneous catalysis. Criteria for useful catalysts are good
selectivity, good activity and good stability. Some examples of
licenced catalysts are:
-Raney nickel, suspended in tetradecahydroanthracene, for liquid
phase dehydrogenation [27J. Provides a yield of 99.6% of MEK at a
temperature of 142°C. Disadvantages are the large amount of
tetradecahydroanthracene (27 times the amount of SBA) required and
the slow convers ion (1.1 kg MEK per kg catalyst per hour).
-ZnO with Bi z0 3 [28J or Na Z C0 3 [29J,supported on brass or steel.
Provides yields of 58 up to 98% of MEK at temperatures between 400
and 500°C. Feed rates are between 1.5 and 6.0 volumes of (liquid)
SBA per volume catalyst per hour. A catalyst example is reported
that af ter 180 days of operation still converted more than 80% of
the SBA to MEK. Catalysts are irreversible poisoned by traces of
water in the feed.
29
-Cu with CrZ03 and MgO on SiO z [30J. Provides at 260°C a product
with 90% MEK, 5% SBA and 5% high-boiling byproducts. Adding 10 vol
% water to the feed provides 95% MEK, 4.8% SBA and 0.2% byproducts.
Reported activity is stabIe over 6 months.
-Copper-tetramine complex with 0.37% CrZ03 [31J. Provides a yield
of 93 to 96% of MEK at a temperature of 270 to 320°C. Low conver
sion rate « 1 vol. liq. SBA per vol. cat. per hour). Regenerated
with air at 350°C and hydrogen at 250°C.
-Cu with BaCrO., CrZ03 and NazO on SiO z [32J. Provides a yield of
97.8% of MEK at a temperature of 180°C. Catalyst is also able to
convert di-secondary butyl ether to MEK.
-ZnO with 6 wt-% CeOz,ZrOz or ThO z [33J. Moderate reaction rate (up
to 6 vol. liq. SBA per vol. cat. per hour), and 1 to 14 mol-% heavy
by-products formed. Maximum MEK yield about 96% at 400°C, but
rapidly decreasing activity af ter 20 hours of use.
3.1.4 Kinetics of a Cu/Ni-catalyst
The kinetics of dehydrogenation of SBA over a catalyst with
composition Cu:Ni:KzO:SiO z (13.8:5.8:0.4:80) have been studied by
Chanda and Mukherjee [18J. Properties of this catalyst are men
tioned in table (3-1):
. BET surracc area (S.) Size Average diameter (d,,) Hulk density (Ph) Pore volume (V.) Porosity (~') Average pore radius (r) ParticIe bulk density (p,,)
table (3-1): catalyst properties
154.9 m~/g - 48 + 65 Tyler mesh 0.02515 cm 0.7188 g/cm 3
0.4519 cm3/g 0.38 58.35 x I O-R cm 1.160 g/cm 3
Analysis of their data shows that a mechanism of dual-site
surface reaction is applicable over the entire temperature range
studied (250-310 0 C).
Below 250°C the conversion was found to be very low while above
320°C the convers ion was found to decrease with increasing
30
temperature. This was due to fouling of the catalyst by reaction
products formed at elevated temperatures. However, in the tempera
ture range of 250°C up to 310°C the dehydrogenation reaction was
not accompanied by any side-reaction and no byproducts were
detected in the reactor effluent. The catalyst which has been used
at 320°C and above regained more than its original activity af ter
it was oxidized with air at 350°C. Stability tests showed no
decrease in activity over a long period of time. It is, however,
recommended to do supplementary tests to make sure the catalyst
keeps sufficiënt activity over a period of two years when it is
only regenerated in the reactor with air at 350°C when necessary.
Other experiments, which were conducted with catalysts of par
ticle sizes in the range of 0.25-1.0 mm diameter (d ), showed that p the rate of reaction remained constant for particle sizes below 0.5
mm, thus indicating the absence of internal diffusional resistance
below this size.
The initial reaction (p(H&) = p(MEK) = 0)
SBA ------) MEK + H&
is a first order reaction with respect to the partial pressure of
SBA. The initial reaction rate ro can be fitted to an equation of
the form:
The values
mentioned
ro = ko * p(SBA) (21)
of the rate constant ko for several temperatures are
in table (3-2), together with the values for the activa-
tion energy.
31
table (3-2): initial reaction rate constant ko at
various temperatures.
temperature (Oe)
250
260
270
290
310
Activation energy: 21.96 kj/mol
ko (mol/g.hr.atm)
0.6279
0.7560
0.9340
1.1180
1.2830
The reaction mechanism of the equilibrium reaction
SBA
K p
) MEK + Hz i-(----
is one of a dual-site mechanism, with the adsorption of alcohol as
rate limiting step. The reaction rate r is derived from the
equation:
ko * (p(SBA)
r = 1
( r in mol ) g.hr.atm
p(MEK) * p(Hz)
K p )
(22)
In the temperature range from 270°C to 310°C the k-values are
given by (T in K):
k H 2.70 * 10-3* exp( 3.92 * 10 3
) = T (23)
kM 0.226 * exp( 0.87 * 10 3 ) = T
(24)
k MH 5.25 10-14* exp( 15.74 * 10 3
= * ) T (25)
32
3.1.5 Pressure influences
From eq.(22) it is obvious that with increasing SBA pressure the
reaction rate also increases while with increasing MEK and Hz
pressure the reaction rate decreases and the equilibrium changes in
favor of SBA. In a tubular plug flow reactor a high pressure drop
over the catalyst bed would be useful for a fast initial reaction
rate (p(SBA) high and p(MEK) and p(Hz) both low) at the entrance of
the reactor and a high degree of convers ion at the end of the
reactor (low total pressure, in favor for equilibrium). This
desired pressure drop can be obtained wether by high flow rat es
(disadvantage: short contact time, so large amounts of catalyst are
required or large SBA recyle will occur) or by the use of catalyst
particles with small diameter (advantage: no diffusional resistance
limitations, resulting in efficiënt use of catalyst area).
The pressure drop over the reactor is calculated, using the
Ergun-relation for the pressure drop over a bed of spherical par
ticles for turbulent gas flow (Re> 700):
Ap
with: E -p -
u -g H -
d -p
u z g
voidfraction
density of gas
gas velocity
height of bed
diameter of particles
H * -a-p
(-) (kg/m 3
)
(m/s)
(m)
(m)
(26)
(the lowest Re-number is later on determined as 1382, what jus
tifies the assumption of turbulent gas flow).
Pressure and pressure drop in the reactorbed are related to the
degree of convers ion of SBA in the bed, because with proceeding
conversion the total gas flow rate increases (one mole of SBA is
replaced by two moles of product). The reaction rate at an ar-
bitrary place in the reactor,
pressures of SBA, MEK and Hz.
however, depends on the partial
A small computer program is written to make an accurate estima-
tion of the expected pressure drop and convers ion in a reactor
tube, filled with catalyst particles. Therefore the tube is cut
into a great number of slices. In each slice the pressure drop is
33
calculated, assuming the SBA convers ion in the slice not having any
affect on the total gas flow rate. At the same time the convers ion
is calculated, assuming the pressure to be constant in the small
slice. Both gas flow rate and gas composition are then adjusted and
used to calculate the pressure drop and the convers ion in the next
slice. Main variables in the program are the initial gas flow and
composition and the initial pressure. Tube length and particle
diameter have fixed values. The output of the program contains,
among other things, the final pressure (must be slightly above
atmospheric) and the degree of SBA convers ion (must be above 90%).
Satisfying initial pressures and flow rates are found by trial and
error. Af ter that, changing the number of slices then gives an idea
of the obtained accuracy. The program has been written in Turbo
Pascal and is to be used on a personal computer.The listing is
presented in appendix (A-5).
3.1. 6 Design
For sufficiënt heat transfer relatively small reactor tubes are
choosen (diameter 0.10 mand height 0.85 m). Each tube is filled
with 4.800 kg catalyst and the maximum initial flow rate with which
a convers ion of 90%, at a temperature of 310°C, is reached, is 0.71
mol/s (189.4 kg/hr). This implies a convers ion rate of 35.5 kg SBA
per kg catalyst per hour. The initial pressure is 2.4 atm. To give
an idea about the catalyst capacity, increasing the initial SBA
flow to 1.42 mol/s
convers ion of 85.6%
and the initial pressure to 4.4 atm, gives a
and a conversion rate of 67.2 kg SBA per kg
catalyst per hour. In appendix A-5 is also the program output
listed for these two cases. In the first case an amount of energy
of 23.84 kW must be added to the tube and in the second case 45.23
kW. With a total initial flow of 1.478 kg/s 99.8 wt-% SBA, an
amount of 28 reactor tubes, each with a length of 0.85 mand a
diameter of 0.10 m is required. The total heat flow from the fur
nace to the tubes must be 89.28 kW/m z tube area.
The minimum required wall thickess t of a reactor tube is found w by the expression [41]:
t w
with: R - external tube radius
p - pressure difference
over tube wall
34
(27)
(m)
(bar)
S - allowable metal stress (bar)
For special Cr-Si-Mo alloys, used in furnaces, the factor S has
a value between 440 bar and 1220 bar. With S = 440 bar, p = 2 bar
and R = 0.10 m, t becomes 0.4 mmo w
With an initial SBA flow of 0 . 71 mol/s the required energy in
the first fifth part of the reactor tube is 14.84 kW or 278 kW/m 2 •
With a thermal conductivity of 17 W/m.oC (average for special
alloys) and a wallthickness of 2 mm, the 6T over that part of the
tube wall must be at least 33°C and the temperature on the outside
of the tube 343°C. This is not a problem in a furnace, where at
800°C, heat transfer is for about 80% obtained from radiation and
for only about 20% from convection.
3.2 Hydrogen recovery
The next threatment for the gas leaving the reactor is to cool
it down and to liquify the major product. At first heat is
recovered in a heat exchanger where the effluent is cooled from
3l0oC down to 210°C and the feed is heated from 99.5°C to 197°C.
Af ter that, the effluent is cooled down to 80.5°C, the required
feed temperature for the first MEK purification column. At this
temperature MEK and SBA are condensed and separated from the
hydrogen in a gas-liquid separator. The remaining gas flow is
cooled further in two stages to remove the remaining SBA and MEK.
In the first stage it is cooled to 40°C with normal cooling water
and in the second stage it is cooled to -5°C with freon. At -5°C
the vapor pressures of SBA and MEK are respectivily 97 Pa and 340
Pa. The hydrogen can therefore be withdrawn at a temperature of -
5°C and approximately atmospheric pressure with a purity of 99.6
vol-%. If the hydrogen is to be obtained with a higher purity,
further cooling will not have much effect and it is better to wash
the hydrogen with a high boiling solvent.
3.3 MEK purification unit
The components in the process stream which have to be separated
are MEK, SBA and a trace of water. The trace of water made it very
difficult to separate SBA and MEK in one column. Simulations with
35
_. __ ._- - ------------------------
PROCESS with the binary system MEK-SBA gave no major problems. MEK
could be separated and obtained with a purity exceeding 99 mol-%
and a yield of 94 % in the top of a column with 30 equilibrium
stages and a reflux ratio of 3. Adding a trace of water (0.5 mol-%)
to the system made the MEK yield decrease to 51.8 %. Increasing the
number of stages and the reflux ratio showed only marginal
improvements. All the water was found in the top of the column,
which means that the bottom only contained a binary SBA-MEK
mixture. In a second column this mixture could easily be separated.
Therefore two columns we re simulated. Figure (3-2) shows the com
position profile of the three components over the first column
(T43) and figure (3-3) does the same for the two components in the
second column (T5l).
36
Mole fraction 1.B
B.5
B.B 1 5 1B
figure (3-2) : composition profile for
(stage 1 is in the top)
MEK purification column T43:
Number of stages
Reflux ratio
Feed : at stage
temperature
pressure
composition: MEK
SBA
Hz.O
Top: rate, relative to feed rate
temperature
pressure
composition: MEK
SBA
Hz,O
Bottom: rate, relative to feed rate
temperature
pressure
composition: MEK
SBA
Hz.O
37
~ HZO
15 ZB tray no.
column T43
20
3
7
80.41 oe 1. 06 bar
89.46 mol-%
10.04 mol-%
0.50 mol-%
45.05 mol-%
78.31 oe 1. 00 bar
98.15 mol-%
0.73 mol-%
1.11 mol-%
54.95 mol-%
87.65 oe 1.19 bar
82.34 mol-%
17.66 mo1-%
0.00 mol-%
.. ale fraction 1.8
8.5
8.8 1 5 18 15 28 25
tray no .
figure (3-3): composition profile for column T5l
(stage 1 is in the top)
MEK purification column T51:
Number of stages 25 Reflux ratio 3
Feed: at stage 7
temperature 82.18 oe pressure 1. 06 bar
composition: MEK 82.34 mol-%
SBA 17.66 mol-%
HzO 0.00 mol-% Top: rate, relative to feed rate 82.73 mol-%
temperature 79.38 oe pressure 1. 00 bar
composition : MEK 99.31 mol-%
SBA 0.69 mol-% Bottom: rate, relative to feed rate 17.27 mol-%
temperature 79.38 oe pressure 1. 24 bar
composition: MEK 1. 03 mol-%
SBA 98.97 mol-%
38
IN Voor-waarts
M Q M
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Massa in kg/s Warmte in kW
Fabri eks voorontwerp No: 2.6Cj3
12.0.~f
45· ? çs
020 -t ~~,~ ~G la
~l~~~QLJ +
l~gljl.S
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S~l ç\).'f~c. ~
Wct.'e'(' ~u.. \'" \ ~û. \ ~oJl
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Totaal:
Apparaatstoom
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Su. \ ~ÛX\ (' ("'Ar ~(ll Wf'l:\,~\",
~v..~ \Su..\ ~oJe v
~(' ~\L'-'\ t'1. \('I\\\'o \ v
Totaal:
M in kg/s Q in kW
1
M Q
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1.02.01- 1'9.512-
6 M Q
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f Compo'nenten .
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'S u. W"" 'f\c.. ö...tÀ.d wnlpr"'
Stt.. B\.thl areola I Sc.,..\:,l..,. U. ,1,..", ',la .,
Totaal:
M in kg/s
11
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Totaal:
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Totaal:
M in kg/s Q in kW
2\
M Q o .qgs I 1'3· 1-9 ,.341- 386'·51
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26
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-
Totaal:
M in kg/s ct in kW
33
M a O·ou \ 8 0'1 S I I. 4:;~ 411.6 <6 0.0 0 1'2- Loo 6
',4=1 ~ Lf';}4.44
3~
M U 0'\:>0\3 O· 61-(). 'l.../ 9 37· '6 r ,. 2l) I 2'$.8' O'o'lb 41, t q
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40 41 42 43 -
M U " M U M Cl. M Q
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A pparca t sTr oom LIet 45 4(, Lfi L(~
t Compo"nenten M Q M a. M a. M a. M a Wcd·e.~ 0.005 '.64{, 0.0\3 \'6 o·s 82 0 ' 00\8 o.'So2...
.S!c. _ Bu.~\ n\Cn'-"n\ 0.0/37- 3.01Lf o.()oLi9 1.0'99 o.()oQ9 o . S 63 o. (44 '36". 311 o. (49 43. IS ):1~\h~ ~ ..... \ Y'.-\. ....... , /.-:;.8:J 2/U,·4g 0.6 3 g (0'2. .113 0·63 g 46.Q?f Cl· 653 '2 q. q, o-oUIZ ,.00' I
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I I
,
-
Totaal: 2.LJ6~ IS 20 ' 1- 2..468 '3/11. ':f 0'656 l 04,2. 0.656 QS .'2g
r-.1 in kg/s Stroom /Comr n in kW ----------------------------------------------
-- - --_._ - ------ -_._---- -
Apparatenli~st voor warmtewisselaars, fornuizen -----------------------------------------------
Apparaat No: H 2 ft q H \3 \-\ \~ \-\ \ ~
Benaming, C O"Y\.cl e-.t\S<I\. {oo Ier \+eaJ-t/L Cooler H eateJL type
Medium
pijpen-/ Ccl ci~·C;o\u..t~", HzS 04-( Ul.~oc{ + ~ a\ . ~ a(ec./ NC1.0HI mantelzijde V ,,_10 \A..\-<'V\e CCA.. Jl. SO \~\:, ~ ( S~e.o..~ C. WÛ\UL CS \ (D.W"\
Capaciteit,
uitgewisselde Lf~<j.63 5 0 '3.'6 1305 Iq \g~(\.~ '2-6. ~ warmte in kW.
Warmtewisselend 26./ oppevl. in m 2 21.0 /4.6 f,~ o. So
Aantal pafKlfê{ 1 1 1 1 1 Abs. of eff. * druk in bar
I pijpen- / 1/3 3/1 \/3 \ / \ l I ~ mantelzijde
I temp. in / uit , . or.
~n ,
pijpzijde '2.0 /40 \4b / 25 52 / ~ \ ~ \ I gg.!" 25 /~t mantelzijde tt /25 20 /40 l~o 1/30 20 I qo J~o 1130
Speciaal te ge- I?. V s 3 \~ {(\JS1lb e.VS 'S\b bruiken mat.
i( aan~even wat bedoeld wordt
Apparaat No: H 211 H 2'1- 1-1 30 H 32 H ~S
Benaming, ~e6<.:> ~ tUl. [ ()Y\.c1 ~eIl ~e \ooae.n.. \\ e.",-t
Co ol ClL type
.ti C-~eJ'\.
Medium b.u:l. cJ.c.uU / c. wo..teAl WaJeJL / ~.~)J W. J~".( pijpen-/ b~.a.O~
C. vJ~ mantelzijde S\-eo.W"\ -+ ei .. ,·su h.:i:L <t> \- ea. 0,1' I\Ebz ,-Capaciteit,
uitgewisselde 220 \ 0:$ 0 warmte in kW.
Warmte .... isselend LI. Q S·S 2 oppevl. in m
Aantal paf!lt~{ \ \ \ I , Abs. of eff.
i(
druk in bar
pijpen- 1 l /3 I I r 1 /3 \ I \ \ I , mantelzijde
temp. in / uit . oe ln ,
pijpzijde ~9.r / 9Q·f' 2.:> 14.:; /00 I (cv CJ~.s Il/o 99.1 11 =1-9 mantelzijde 130 1IJu ~:;.s I:;:;· r 11.)IUo ~ IC/u '3/0 12/;;)
Speciaal te ge-bruiken mat.
7( aaneeven .... at bedoeld .... ordt
Apparaat No: r- 36 ~ 3CO H ~o \-\ q, \-\- 4 <J
Benaming, ç'~U- C~~ Coolen- Coo \-eJL eebo',\u type
Medium ~~~. cJc.t.e. / c. watiA; c. (.AJ~I e~~/ ~~CJ.A el ... \0\. pijpen-I
ç'~ ~J 5J-~), + HEt H~cl(,,~ I s+eo..~ mantelzijde t\'do.(O~
Capaciteit,
uitgewisselde 2.25 ~o \. 6 2..0 ,91.. 2.0 . 0 842.. S 'Warmte in kW.
Wa=mte .... isselend /, '13 o. ~ 1\, ~ oppevl. in m 2 0·5
Aantal pafältê{ 1 1 1 1 1 Abs. of eff.
i(
druk in bar
pijpen- I 2·'-411 1 11 1 / 1 1 ( 1 \ (1 mantelzijde
temp. in I uit . °c l.n ,
pijpzijde \ '1'1- I 3t.:> 2.:) 14 û LO 11.4 0 -5/-~ ?>B{8~ mantelzijde Soo/)JI.:l 210 ( 80· S" 30. r I Cf û ~o {-5 130 I 1J0
Speciaal te ge-bruiken mat.
I
i( aan~even wat bedoeld .... ordt
55
Apparaat No: ~I Ll t- t-t 52 H S5 H 51- H S~
Benaming, C~eY\SVL ~Ioo~ \m. C~~~ Cao 1 efL Cao lOL type
Medium
hJ· alcoh~ C. w~1 r\E.~ r\E~ c. wo1(/L! pijpen-I
mantelzijde I1E~ SteAW"\ MEK C. woleIL C. WcJvt
Capaciteit,
uitgewisselde <O~\·b 1ll5. 0 112..'3, 59 '56 . 1-'Warmte in kW.
Warmte'Wisselend (Ö'b L-6 \ L{. '3 \ . Ll oppevl. in m 2 I ,Lf
Aantal pafäItê{ 1 1 1 1 ~ Abs. of eff. ~
druk in bar
pijpen- I 1 I 1 mantelzijde 1 I 3 '1 11 1 l1 1 (1
temp. in I uit . o~ ~n ,
pijpzijde '20(4° 105 1(0) ?a 140 -::t-qll-J o 15 / q0 mantelzijde 16/1-3 '5 0 /'Jv -=1-9 l1-ct Lo/q." LO l q lJ
Speciaal te ge-bruiken mat.
i{ aan~even 'Wat bedoe~d wordt
_ ._------ .... __ ._-
Technische Hogeschool Delft Afd. Chemische Technologie
Fabrieksvoorontwerp No: .2' ~ 3. . Datum : • • g 11- .{ Iq·S t6. . . . Ontworpen door •. A.\-\: l.W\(.'ç' •
Y?~E!11:~~Iê.§ELMgsP~fI!:I~1:I~~~~ R, JA- (~~\-V'-Apparaatnunnner : H. 2 . Aantal : .1. serie/parallel *
ALGEMENE EIGENSCHAPPEN :
Funktie . . • . • . . • • :
Type . . . . . . . . . . :
Uitvoering . . . . . . . :
Positie . . . . :
TT '~ * dtaiii'is keselaar iC. alG!! Kondensor Uuie.,,,,
met vaste pijpplaten* ~, • B J 1
1 ... 11' .1j bhh 2 pj jp p2 'iuu e_'ililFkllllslaar
horizon taal/'u T'i Lh ur*.
Kapaciteit
Warmtewisselend oppervlak : .. 4 ~'\~ '1-: ... 7..-6. \ ..
.kW (berekend) 2 • m (berekend)
Overallwarmteoverdrachtscoëfficiënt : .. 56S". 2 • Wim K(globaal) Logaritmisch temperatuurverschil (LMTD) • : • • Aantal passages pijpzijde •••••••• : 4 Aantal passages mantelzijde • • • • • • • : 2
14.3 . .. Oe
Korrektiefaktor LMTD (min. 0,75). : o .. tO o Gekorrigeerde LMTD. . • . . : .. \~.H!i.
BEDRIJFSKONDITIES :
Soort fluidum . .
Massastroom . • . kg/s Massastroom te ~~l~I .. /kondenseren~ . • • • • • .kg/s Gemiddelde soortelijke warmte . Verdampingswarmte • •
Temperatuur IN
Temperatuur UIT .
• • • • .kJ/kg •. oe
. • • • • . • kJ/kg
° • • C Druk . . . . . . . . . . . . . . . . . . . . . . . bar Materiaal .
o •• C
Mantelzijde
t'l .. b. t;~ •
· .\~o.q .. ..\.Oc;.t .. · . ~. ~ 5 . . · . '!SS - . · ·11· . . .
· . ·3· . . . . · .S~uJ .. .
Pijpzijde
Ca. de ~ot~\~", · be'L'O .. · .. ~ . • ~.a 'i) •
· . 2-.0. . . · . . 4-a .. · . . 1 . . . · .~~~ " .
*Doorstrepen wat niet van toepassing ~s
·57 -------------------------
Apparatenlijst voor reaktoren, kolommen, vaten ----------------------------------------------~----------~----------~----------r_--------~----------~--------~ I
Apparaat No:
Benaming,
type
Abs.of eff.*
druk in bar
T3 V6
1 1 1 ~-----------+----------~----------~--------~----------~--------~ '
temp. in oe
Inhoud in m3
Diam. in m
1 of h in m
25-40
2. 2. \ 1·5
12·5
l46 liO
0 •• 64 o. blo (l,bO
\ 11
\.00
3.1-\e ' f
52.. I
0·150
0.50
0·1-66 ~-----------+----------~----------~---------+----------~--------~ I
;{ Vulling:
schotels-aant.
vaste pakking
katalysator-
type
- ,t - vorm
Spec iaa2. te ge-
bruiken ::lat.
aantal
serie / :;:a:ëallel
CL LIS 1\(,
1
* aangeven wat bedoeld wordt
Apparatenlijst voor reaktoren, kolommen, vaten ---------------------~------------------------
Apparaat No: T tl{ t"\ tb T 17- \I 20 T 2.3
f-\c.;d No..o t\ A ~s\)( r\~0\'1 Li,\~ - [) \ ~\: \~;J~~ Benaming, sh·:p~.ljL L\',~ type
~\b"~ ~lu.~>') löl\A""'" '1 I
5..e.f~óL 3(
Abs.of eff. 1 1 1 1 1
druk in bar
temp. in oe '3\ - 10 'l ~\ <j\ gfJ· 5 '.f.1. f - \Ct)
Inhoud in m3 ~.S "\ +.~S 1,08$ 2 \. L Diam. in m \. \ 0 0·20 1 . 0 0,7° \. S 1 of h in m \0 1..0 1 0 2,80 \'2. Vulling:
3(
schotels-aant. \3 TMTP- ss 23
vaste pakking i>~~~ katalysator-
ItoDM;-'.1 oll~-type ~~ 2.5 he,...
- " - vorm
· ............ · ............ · ............
Speciaa.l te ge-R.. V S 1 tb f- V5 316 e.. V j J { l.
b::uiken :nat.
aantal 1 1 1 1 1 serie/ pa!'allel
* aangeven wat bedoeld wordt
53
Apparaat No: '-J 2t; \" '2-'1 \J ~~ ((.3'1- Cj 3" S~.
L\~~ - O\I)\~ \\~~~~ O!~ ~~-Benaming, ~6L type \..t\~ Co \ u.. "'" "" S-\<5\o8f- Lt~~
~()j\J.ó\- ~~óL 3(
Abs.of effe 1 1 1 2,Q 1 druk in bar
temp. o 0c ~n 1?-- ~ =1-7- r - loa 4-0 '2> \ 0 30.)
in m3 2'b 1. (J 2-0 " Inhoud ,01 0-0
Diam. in m /'" l o-s" 3 0 25 0·25'
1 of h in m ~.'2..S la '2. ~1 .~ l . 2 S- I
Vulling: :;{
schotels-aant •. \3 vaste pakking
katalysator- C"'/,{{ Oh );0, type S~~J
- " - vorm
· ............. · ...... . ....... · ...... ~ .........
Speciaal te ge- CÎl,-)l--Ho-b:-uiken ::lat. a/lo}"
aantal 1 1 serie ,lDa!'allel 1 1 1
3( aangeven wat bedoeld wordt
60
Apparaat No: T 4-3 V LIl, T5\ " 54
Benaming, O;St-; llo..t· o~
UtSsJ DiS\; \\o..\ttM
\J f! '>S~ type Co lu\'Y'\ Y'\ [Qtûw\.t) !
I 3{ I Abs.of eff. 1 ~ t \ I druk in bar
I temp. in oe -:tg - 2,<6 tg '1-l3 - l" S 1-9 I
in m3 I
Inhoud ll, ~ , 15·0 , I
Diam. in m 1 , I I I
1 of h in m 15 1 • 2. 19 I·Z I
Vulling: 2(
schotels-aant. 36 2~ vaste pakking
katalysator-
type
- , , - vorm
· ............ · ............ · ..... . ......
Speciaal te ge-
b::-uiken :':lat.
aantal 1 \ \ ,
se::::,ie / pa::::,al lel
1( aangeven wat bedoeld wordt
Technische Hogeschool Delft Afd.Chemische Technologie
Fabrieksvoorontwerp No: 1. b ~ '1 Datum: 1- /1- , \ Cj '6 fb Ontworpen door : ' A. \-\. ~"1"'\~
TORENSPECIFIKATIEBLAD R.. c4 f"~ \ tA-
Apparaatnummer : T··5 I Fabrieksnummer :
ALGEMENE EIGENSCHAPPEN :
Funktie ............... : destillatie / a_,liiiiiil~iliiii , i~iiliiiii~1: / •••••••• ;te Type toren ... 0 •••••••• : 8 R \ultE / schotel / iipt:geilu~ / * ................. Type schotel .......... k1.91-j Oi / zeefplaat / Hiàlue- / * : .................. Aantal schotels ....... : theoretisch : "2..0 Aantal schotels . ...... : praktisch : '2..~ Schotelafstand / HETS : o· S· m Materiaal schotel : ~~~ Diamet er toren ........ : \ • .0 0 m Hoogte toren .... : \~ W"\
Materiaal toren ...... : s-\ui Verwarming ............ : ~/ 'ilfilOiiil ij e 11111" / reboiler / * ................
BEDRIJFSKONDITIES :
Voeding , Top Bodem Reflux/-eb~ol:p E~tt'e\tüQ ~e Ifti:èèe:l .miàèei/ •.•
Temperatuur ...•. oe 82> *1-9 \05 :tij Druk .. 0 ••••••••• bar , . \Cj I \, '2.4 \ Dichtheid ....... kg/m 3
5 0 1- <005 C6lS 80S Massastroom ..... kg/s o . :}Cf t o ·6Sb o· \ S \ . .g \ Samens telling ln "'" ,,/. VI~. W\ '10 v'l. _.,. w,I. W\. ". w ·10 mol 7. resp. gew.7.
Se-.c.o~o\ \1·1- \~ oob~ Oot ~8'~1 '1 Cf" J ,,·bq 0':;
M EK g'2·3 82 ~Oj,J' 9~·J /,0 J 0,9 '\'1.J I Cf'" 3
ONTWERP :
Aantal kj J ' / zeef gaten / ** lo.fo T)l'5e flekkittg ...•.. : ... :
Aktief schoteloppervlak .... o.bj m 2
Hete!':tee:3: l' e:td~ iftg - . ..... : : Lengte overlooprand ......... ..... : ~oomm AftHetirrgcn pakking : Diameter valpijp /~ / ............. :
7 I!.\W\HR
Verdere gegevens op schets vermelden
* doorstrepen wat niet van toepassing lSo
! i
**, d " d- ,
ln len een toren schotels van verschillend oncwerp bevatJ .~l~t~a~alin~g~e~v~e~nL. ______________ __
Apparatenlijst voor pompen, blowers, kompressoren --~----------------------------------------------
Apparaat No: c1 eS p \0 r 11- PIS I
Benaming, CCJWlfrt.,sS 0'- ~"~'JJ c4~j ~~~ ~~Jl
type f(Á~f p~p fu.~p f tA.W"\f
te verpompen &>0 wt_o~ 8a wt-% Hz SOl\-
fl_ b \A.. t (..vU. H2.S o~ HL 30y No..o~ I medium + W.rJCb~ I
Capaciteit in \.oLt '2,2- S 6.25 1. 2 1- 0, lL~ I
t/d of kg/si( I
2.'50..\/ l I
Dichtheid \~OO ( 2 l q I 15 Z l 0 0 u I
kg/m 3 ~ in (./:r.e 1;:, ZI '" •
~l~~;p 1
in bar(abs. · 3.2 S I ~i()
0
ltf6 [ lqb lol.! (01.. temp. in C 25 / 1-1 5 L I 5 2 <3 \ I ~ 1 in / uit I
I Vermogen in kW 13. +/ theor./ prakt.
\ 0':>.3 Speciaal te ge
R- tts bruiken mat. 5l~ eVS J I , f..VS ~\b e.rlS 3l€
aantal 1 serie / parallel 1 1 1 t
* aangeven wat bedoeld wordt
Apparatenlijst voor nomuen. blowers. kompressoren -------------------------------------------------
Apparaat No: P '2.. 1 P '2. 2. P 'LS P 2 <ó r ~ \
~~j) ~~\\.\.~
~J) cdf,~ Benaming, ~j~ ~7JJJ type
t>\.L~f fu.~ f~p ~ ~
B~. Jc.JJ. gvl,- c..O(M ~~- cJea1J B~_ o1cv{) wa.i0-te verpompen + Di; \ u L..",,--:t-vf ....L-+ Oi ~So b~
medium +~~ -+ Wa..tVL -\-wo:ttA
Capaciteit in \·3'1 4 ,6 b C6 o -1-5 Cf o -s I t/d of kg/s~ \. g 1
Dichtheid g1- 0 3\5 131> ÛJ1t loo 0 kg/m 3 in
Zuig-/persdruk
1/1.5 . in bar(abs.~
~!) I
0
B~-) (g~- ) Cj~.r / c,S, r ~1,r Itf') 1- f~ s 19-t-;- ! 1 o~ temp. in C 10.;) in / uit I
I Vermogen in kW
O· \~~ theor./ prakt. 0·2 Speciaal te ge I
bruiken mat I
I
aantal
I 1 \ \ \ serie / parallel I
I
* aangeven wat bedoeld wordt
~EE~~~~~~!!~~~_~~~~_E~~E~~~_~!~~~~~~_~~~E~~~~~~~~
Apparaat No: p ~q P 42- P ~ 5 P q~ ~ L\ C'j -
~ ~\v..'Jl . ~J ~J J~.J Benaming,
~J ~j type f7 pc.\.~p r~p (?u -(J H f~
te verpompen ~~.~ M(~ \1 Et:..
t---\ Et: rtE K medium + o.f.û!'
+ d1c.-~ ~ cJc~)
Capaciteit in \ 131 , ,LJQ2 I. \ 6 o·6Qr 6.g t/d of kg/s*'
Dichtheid 3\5 ~o~ ~a ~ ~oS- ~o<Ó kg/m 3 in
Zuig-/persdruk
in bar(abs.of
eff. *.)
temp. in °c qo (C(o So.)"{ 2a,~ 1-8(1-~ 7-2 (1~ ~~ I F~ in / uit
Vermogen in kW I
theor./ prakt. I
Speciaal te ge I
bruiken mat
aantal 1 \
, serie / parallel \ ,
I
I
* aangeven wat bedoeld wordt
p So p'S1 P5b I
Apparaat No:
Benaming , ~~jJ ~ ~J ~7J type fUY>'\f ftAf ~
te verpompen B J- -cJc.o~ f5~ +Á.{J nE~ medium .J... t-\E~
Capaciteit in o· 15 \ , g \ \ a·66 t/d of kg/si(
Dichtheid SiS 30g ~oS-kg/m 3 in
ZUig- / persdruk
in bar(abs.of
eff.i()
temp. in °c loS f{Ö) ?-q t 1-Cf 1-9 {1-cr in / uit
Vermogen in kW
theor. / prakt.
Speciaal te ge
brui ken mat
aantal 1 serie / parallel \ \
* aangeven wat bedoeld wordt
66
6 Cost estimation and economics
The equipment costs have been calculated, according to the
prices published by the Dutch Association for Chemical Engineers
[24] and are given in the following tabie:
table (6-1): Equipment Costs
Equipment Number
Compressor and 1
cooler 1
Heat Exchangers 18
Columns 7
Tanks and
separators 11
Pumps 17
Furnace 1
Reactor 1
Total 57
Costs (f)
~ 384,000
880,000
443,400
329,000
42,000
126,000
2,654,400
The total capital investment has been calculated by using the
method of H. C. Bauman [23] and is given in the following tabie.
table (6-2): Total capital investment
Component Ratio
Purchased equipment 100
Equipment installation 47
Instrumentation 18
Piping 66
Electrical 11
Buildings 18
Yard improvements 10
Service facilities 70
Land 6
Total direct costs 346
Engineering and supervision 33
Construction expenses 41
Construction's fee 21
Contingency 42
Fixed-capital investment 483
Working capital 86
Total capital investment 569
68
Cost (f)
2,654,400
1,247,568
477,792
1,751,904
291,984
477,792
265,440
1,858,080
159,264
9,184,224
875,952
1,088,304
557,424
1,114,848
12,820,725
2,282,784
15,103,536
table (6-3): Raw material costs
Material
n-Butene
Sulfurie acid
Sodium hydroxide
Total
Amount
(t/yr)
Cost
(f/t)
26,456.6 638
290.3 150
300.9 500
table (6-4): Selling prices
Material
MEK
Hydrogen
SBA
Total
Amount Price
(t/yr) (f /t)
33,721.9 1750*
933.2 1800
o 1550
Total cost
(f/yr)
16,879,311
75,996
,150,467
17,105,724
Total price
(f/yr)
59,013,325
1,679,616
o
60,692,941
*: Price obtained from Shell Nederland Chemie section Marketing.
A price of f 1750/ton was reported as normal although in
june 1988 prices increased to f 2100/ton and incidental
prices up to f 3300/ton where reported.
table (6-5): Manufacturing costs
Component
A.Direct production costs
(60% of total product costs)
1.Raw materials
Ratio
60
(10-50% of p.c.) 33
2.0perating labor
(10-20% of p.c.) 10
3.Direct supervisory
(10-25% of operating labor) 1.5
4.Utilities
(10-20% of p. c.)
5.Maintenance and repairs
(2-10% of fixed capital)
6.0perating supplies
(0.5-1% of fixed capital)
7.Laboratory charges
(10-20% of operating labor)
8.Patents and royalities
(0-6% of p.c.)
B.Fixed charges
(10-20% of product costs)
C.Plant-overhead costs
(5-15% of p.c.)
10
1.5
0.2
1.5
2.3
15
8
Cost (f)
31,101,316
17,105,724
5,183,553
777,533
5,183,553
777,533
96,155
777,533
1,192,212
7,775,329
4,146,842
D.General expenses
1.Administration
(2-5% of p.c.)
2.Distribution and sel1ing
(2-20% of p.c.)
3.Research and development
(5% of p.c.)
4.Financing
(0-7% of total capital)
Total
Income
Gross annual earning
17 8,812,039
3 1,555,066
8 4,146,842
5 2,591,776
1 518,355
100 51,835,527
60,692,941
8,857,414
Two statie methods, used for deciding if a n investment is
economical justified, are the pay-out time calculation and the
return on investment calculation.
The pay-out time (POT) is defined as the minimum required number
of years, necessary to repay the original investment. As is assumed
that the working capital is returned af ter ending the project, the
original investment only consists of the fixed capital investment:
POT = gE~§§_~~~~~l_~~E~!~K ___ _ fixed capital investment (28)
This assumption is not made when calculating the return on
investment (ROl). The ROl is defined as:
ROl = grQ~~_~~~~~l_~~r~!~g ______________________ * 100% fixed capital investment + working capital
(29)
For this project the POT is 1.45 years and the ROl is 58.6%.
The internal rate of return (IRR) is an example of adynamie
method. With this method the cash flows, inc1uding the investments,
over the entire life time of the project are converted to this very
day with a return fraction r. The sum of all converted cash flows
must be zero and this can be obtained by changing r.
7'
For this project the 1ife time is fixed at 10 years and the rest
value RV of the equipment is fixed at 10% of the fixed capita1
investment F. Furthermore it is assumed that the gross annua1
earning E is constant over 10 years. The working capita1 W is
returned af ter 10 years. The converted cash flow over a period of
10 years is:
-F -W + E + E E + E+RV+W (30) Ï+r (Ï+r)Z + ... + (Ï+r)g (Ï+r)ïo
Solving this equation with:
F = 12,820,725
W = 2,282,784
E = 8,857,414
RV = 1,282,073
gives a va1ue for r of 0.58187 and a IRR of 58.2%.
72.
REFERENCES.
1 Kirk-Othmer, "Encyclopedia of Chemical
Technology",vol.13,Wiley Intr.Ed. (1984)
2 G.A.Chernyshkava and D.V.Mushenko, J. Appl.
Chem. (USSR), 53(11), 1834 (1981)
3 Hydrocarbon Processing,48(11),204 (1969)
4 Kirk-Othmer, "Encyclopedia of Chemical Technology",
vo1.4,Wiley Intr.Ed. (1984)
5 C.B.Dale, C.M.Sliepcevich,and R.R.White,
Ind. and Eng. Chemistry, 48(5), 913 (1956)
6 Petroleum Refiner, 36(11), 264 (1957)
7 Petroleum Refiner, 38(11), 272 (1959)
8 Chemical Engineering, Feb. 8, 63 (1960)
9 H.S.Davis ,J. Am. Chem. Soc., 50, 2780 (1928)
10 H.S.Davis and R.Schuler,J. Am. Chem. Soc., 52, 721 (1930)
11 R.C.Weast,"Handbook of Chem. and Physics" ,
The Chem. Rubber Co. (1971-1972)
12 R.H.Perry and C.H.Chilton,"Chemical Engineerings Handbook",
McGraw-Hill Int. Book Co. (1974)
13 J.Racz, J.G.Wassink Hnd P.Dees, Chem. Eng. Tech.,
46(6), 261 (1974)
14 E.E.Ludwig, "Design for Chemical and Petrochemical plants",
vo1.2, Gulf Publishing Co., (1964)
15 I.Rodger, Chem.Eng.Progress, 11(2), 39 (1982)
16 G.M.Smith and E.Mantius, Chem.Eng.Progress, 17(9), 78 (1978)
17 H.J.Kolb and R.L.Burwell jr., J.Am.Chem.Soc., 67, 1084 (1945)
18 M.Chanda and A.Mukherjee, J.Appl.Chem.Biotechnol.,
28, 119 (1978)
19 D.D.Wagman and W.H.Evans, "Selected Values of Chemical
Thermodynamics Properties", National Bureau of
Standards, Washington (1968)
20 "WEBCI/WUBO prijzenboekje", Dutch Association of Cost
Engineers, Leidschendam (1986)
21 J.M.Coulson and J.F.Richardson,"Chemical Engineering",
vol.6, Pergamon Press (1983).
22 L.Ricci, "Separation Techniques 1: Liquid-Liquid Systems"
McGraw-Hill Publ. Co. ,New Vork (1980).
23 M.S.Peters and K.D.Timmerhaus, "Plant Design and Economics
--------- ---
for Chemical Engineers", McGraw-Hill Publ. Co., N.Y. (1958)
24 Ullmanns, Encyklopädie der Technischen Chemie,
4, band 9 (1975)
25 Y.Yamamoto and T.Maruyama, Kagaku Kogaku, 23, 635 (1959)
26 I.N.Bushmakin, A.P.Begetova and K.I.Kuchinskaya,
27
28
29
30
31
32
Sintet.Kauchuk, 4, 8 (1936)
US pat.no. 2,829,165 (1958)
US pat. no. 2,436,970 (1948)
US pat.no. 2,835,706 (1958)
German pat. no. DT 2,347,097
German pat. no. DT 1,026,739
German pat. no. DT 1,913,311
(1973 )
(1958)
(1969 )
33 British pat.no. 663,376 (1949)
34 T.Magnussen, M.L.Michelson and A.Friedenslund,
IChE Symp.Series 56, Proc.Int.Symp.on Dist.,
London (1979)
35 L.M.Naphtali and D.P.Sandholm, AIChE Journal,
17(1), 148 (1971)
36 J.Gmehling, U.Onken and W.Arlt, Vapor-Liquid Equilibrium
Data Co11ection, vol.1 (suppl.1), Dechema (1981)
37 A.G.Montfoort, De Chemische Fabriek, deel IA: Flowsheet
theorie en ontwerp, Collegediktaat TUD, Delft (1980)
38 A.G.Montfoort, De Chemische Fabriek, deel 11: Economische
aspecten en cost-engineering, Collegediktaat TUD,
Delft (1980)
39 F.J.Zuiderweg, Fysische Scheidingsmethoden, Collegediktaat
TUD, Delft (1980)
40 concept diktaat Apparaten voor de Procesindustrie deel 4:
apparaten voor warmteoverdracht, Collegediktaat TUD,
Delft (1980)
41 W.L.Nelson, "Petroleum Refinery Engineering",
fourth edition, McGraw-Hill Publ. Co., N.Y. (1958)
7'1
Appendix
A-I Chemical and physical properties
table (1): Antoine constants
In (p) = p in mm Hg and t
Component A B c
n-butene 15.785 2299.6 -22.77
15.737 2932.1 -52.55
SBA 17.210 3026.0 -86.66
DiiB 18.585 3984.9 -39.73
MEK 16.264 2904.3 -51.19
in K.
:temp. range
°C
: -73 +27
+1 - +100
:+25 - +120
-2 - +127
: +20 - +120
Tab1e(2): Chemica1 and physica1 properties
Methyl Ethyl Ketone
-Molecular weight
-Boiling point at 1 atm.
-Freezing point
-Refractive index,
-Density at 20°C
-Surface tension at 20 0 e
-Specific heat of vapor at 137°e
-Specific heat of 1iquid at 20°C
-Heat of combustion at 25°C and
constant pressure
-Heat of formation at constant
pressure
-Latent heat of vaporization at
79.6°C and I atm.
-Critical temperature
-Critical pressure
-Viscosity at 20°C
-Flash point
-Solubility in water at 20°C
-Electrical conductivity at 20°C
Secondary Butyl Alcohol
-Molecular weight
-Boiling point at 1 atm.
-Freezing point'
-Refractive index,
-Density at 15°C
-Specific heat of vapor at 137°C
-Specific heat of 1iquid at 20°C
-Heat of formation at constant
pressure
-Latent heat of vaporization at
99.5°C and 1 atm.
72.10
79.57
-85.90
1. 378
804.5
24.6
1732
2084
2435
-279.5
32.8
260
4299
0.416
-1
27.5
2*10"
74.10
99.5
-114.7
1.39446
810.9
2730
-268.1
41. 687
°c
°c
kg/m 3
mN/m
J/kg. oe
J /kg. oe
kJ/mole
kJ/mole
kJ/mole
°c
kPa
mPa.s
°c
wt-%
pS/m
°c oe
kg/m 3
J/kg. oe
J/kg. oe
kJ/mole
kJ/mole
-Critical temperature 265 °c -Critical pressure 4850 kPa -Viscosity at 15°C 42.10 mPa.s -Flash point 24.4 °c -Solubi1ity in water at 30°C 18 wt-%
71
n-Butene (2 % l-butene. 89 % trans-2-butene.
9 % cis-2-butene)
-Molecular weight
-Boiling point at 1 atm
-Freezing point -25 -Refractive index, nD
-Density of liquid at 25°C
-Density of gas at OoC and 1 atm
-Surface tension at 20°C
-Specific heat of vapor at 25°C
-Heat of combustion at 25°C and
constant pressure
-Heat of formation at constant
pressure
-Latent heat of vaporization at
lOC and 1 atm.
-Critical temperature
-Critical pressure
Sulfuric acid
-Molecular weight
-Boiling point at 1 atm
-Melting point
-Density of liquid at 25°C
-Surface tension at 20°C
-Specific heat of liquid at 20°C
-Heat of formation at constant
pressure
-Critical temperature
-Critical pressure
56.11
0.99
-110.1
1.3868
602.09
2.591
0.01356
1550
647.1
-9.443
21. 60
155.9
4147
98.08
338
3.0
1.841
50
1443
-811.2
655
8208
°c
°c
kg/m 3
kg/m 3
mN/m
J /kg. °c
kJ/mole
kJ/mole
kJ/mole
°c
kPa
°c
°c
kg/m 3
mN/m
J/kg.oC
kJ/mole
°c
kPa
A-2 Stream data compressor Cl
VERSION 0484 ***********:::
SM PROCESS INPUT LISIING - PAGE 1
GENERAL DATA TITLE USER=A\ AND R ,PROBlEM=CCMP,PROJECl=FABONI,DAIE=FEB87 DIMENSION SI,TEMP=C,PRESS=BAR PRINT \.lTO fIletl
COMPONENT DAIA LIBID 1,BUI1/2,BTC2/3,BTT2
THERMODYNAMIC DATA TYPE SYSTEM=SRK
STREAH DATA PROPERTY STRM=1,TEMP=25,PRESS=1.0,PHASE=V,*
COMP (M)=1,2.0/2,9.0/3,89.0,NOCHECK,RAIE(M)=66.7 UNIT OPERATICt\S
COMPRESSOR UID=Cl,NAME=BUT-CO~PRESSOR,KPRINT FEED 1 PRODUCT L=2 OPER PIN=1.O,POUT=3.2~,PCLY=76,TESl=40 COOLER DP=Q.25,TOUl=25
VERSION C484 SIMULAIION SCIENCES, INC. PROJECT FAEONI PROBLEt1 Cm1P
SM PRCCESS
SCLUTION
PAGE 7 A. Ar~D R FEB87
SUMMARY OF COMPRESSOR/EXPANDER/PUMP/IURBINC UNIIS
1 UNIT Cl , BUI-COHPHESS, IS A CCMPRESSCR *** FEED STREAHS ARE 1 *** LICUID PRODUCT IS SIREAM 2
*** OEERAIING CONDITICNS
TEHPERAIURE, DEG C PRESSURE, BAR ENTHALPY, MM KJ . /HR ENTROPY, KJ j~OLE DEG C MOLE PERCENT LIQUID ADIABAIIC EFFICIENCY, PERCENT POLYTRCPIC EFFICIENCY, PERCENT ISENTROPIC COEFFICIENT, K POLYTROPIC COEFFICIENI, N HEAD,M
ADIABATIC PGLYT:lOPIC ACTUAL
'WORK, KW THECRETICAL POLYTROPIC ACTUAL
COHPONEt\TS
INLET
25.00 1.0000 2.6381
216.3610 0.0000
ISENTRC!?IC
60.82 3.2500 2.8359
216.3510 0.0000
- CC~PCNE~T KVALUES -1 THRU 3 2.6640E+00 2.1374E+00 2.2247E+00
AFTERCCCLER DUTY, MM KJ /HR TEHPERATURE, DEG C PRESSURE, EAR
1 THRU 3 9.9483E-01 7.3218E-01 7.7219E-01
90
OUlLE!
71.01 3.2500 2.9035
219.3471 0.0000
74.56 76.00
1.1121 1.1529
5391.95 5496.42 7232.14
54.98 56.04 73.74
1.7624 25.00
3.0000
VERSION C484 SIMUlAIICN SCIENCES, INC. PROJECI F tWONl PROBlEM Cm·1P
SM PROCESS
SOLUTION
SIREAr1 SUMMARY
STREAM ID. NAtiE PHASE
FROM Ul\IT/TRAY TO UNIT/TRAY
FROM SlREM1
KG MOlS/HR TEMPERAlURE, rEG C PRESSURE, BAR H, MM KJ /HR
M KJ /KG MOlE KJ IKG
MOlE FRACT LIÇUID
M KGS/HR MOLECULAR ~EI(HT
STD lIQ M3/HR
UOP K
DEG API SP GR KGS/M3
REDUCED TEMP REDUCED PRESS ACENTRIC FACTOR **VAPCRl)*
M KGS/HR MOLECULAR ~EIGHT STn LIQ M3/HR SID ~ M3/HR ACTUAL H M3/HR
KGS/fo! M3 z CP,KJ IKG MOL C
**LIQUID** M KGS/HR MOLECULAR ~EIGHI SID lIQ r13/HR ACTUAI GPB
Z
~13/HR ~GS/~3
CP,KJ IKG MOL C
1
VAPOR Ol 0 1/ 0
66.700 25.000 1.000 2.638
39.551 104.913 0.00000
3.742 56.108
6.134 99.965 0.6113
610.0669 12.926
0.695 0.025 0.217
3.742 56.108
6.134 1.495 1.610
2323.841 0.97403
8.9201E+01
0.000 0.000 0.000
0.0000 0.000 0.000
0.00000 O.OOOOE+OO
2
LIQUID 11 0 Ol 0
66.700 25.000
3.000 --1.141
17.107 304.893 1.00000
3.742 56.108 6.134
99.965 0.6113
610.0669 12.926
0.695 0.075 0.217
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 C.COOOE+OO
3.742 56.108 6.134
30.4500 6.916
541.125 0.01255
1.3545E+02
PAGE 10 A. AND F .. FEB87
r;
A-3 Stream data SBA stripper T14
N 0484 SH SIHULATION SCIENCES, INC. PROCESS PAGL 12 PROJECT FAI30NT A AND H PROBLEM ACID SOLUTION Jl\N87
STREAtl COl1PONENT FLOU RAIES - KG 11OLS/HH
STRSAM ID 1 2 J q NArE
PHASE LIQUID 'JAPOR '/APOR LIQUIU
1 WATER 827.6753 164.0000 99.7643 891.9106 2 SBUOH 65.4740 0.0000 65.'1085 0.0655 3 SULFURIC 65.4740 0.0000 0.0000 65.4740
IOTALS 958.6230 164.0000 165.1728 957.4500 IEHPERATU~E, ~EG C 51.4'.)00 190.0000 9:.0728 101.5943 PRESSURE, BAR 1. 0000 3.0000 1.0000 1.0uOO H, MM KJ /HR 4.3971 8.2191 0.9086 7.7576 HOLE FRACT LIQUID 1.0000 0.0000 :J. 0000 1.0000 RECYCLE CONVERGENCE 0.0000 0.0000 0.0000 0.0000
A-4 Data azeotropic distillation unit
SYSTEM H20-SBA-DiiB at 760 mm Hg
NUMBER OF COMPONENTS 3
COMPONENTS 1 H20 2 SBA 3 DiiB
ACTIVITY COEFFICIENT: O=NONE, 4=UNIQUAC 4
UNIQUAC BINARY 0.000
-73.656 331.969
INTERACTION 259.551
0.000 105.468
PARAMETERS 779.367 259.935
0.000
UNIQUAC SURFACE AND VOLUME PARAMETERS 5.616 3.924 0.920 4.920 3.664 1.400
ANTOINE COEFFICIENTS 15.737 2932.149 17.210 3026.030 18.585 3984.922
8s
-52.545 -86.660 -39.734
NUMBER OF STAGES 16
NUMBER OF FEEDS 2
THE STAGE ~T WHICH FEED 1 IS INTRODUCED 8
THE VAPOR FRACTION OF FEED 1 0.000000000000000
COMPONENT FLOW RATES IN FEED 1 26.920000000000000 18.170000000000000
THE STAGE AT WHICH FEED 2 IS INTRODUCED 12
THE VAPOR FRACTION OF FEED 2 0.000000000000000
COMPONENT FLOW RATES IN FEED 2 4.660000000000000 12.990000000000000
CONDENSER ( YIN)? Y
THE DISTILLATE RATE 77.120000000000000
THE REFLUX RATIO 4.200000000000000
NUMBER OF LIQUID SIDE STREAMS o
NUMBER OF VAPOR SIDE STREAMS o
THE PRESSURE 760.000000000000000
ESTIMATE THE TOP AND BOTTOM STAGE TEMPERATURES IN DEGREES CELSIUS
77.000000000000000 99.000000000000000
THE MAXIMUM CHANGE IN TEMPERATURE BETWEEN ITERATIONS ( DEGREES CELSIUS) - OFTEN 10
2.000000000000000
THE MAXIMUM FRACTIONAL CHANGE IN FLOW RATES BETWEEN ITERATIONS - OFTEN 0.5
0.100000000000000
0.0000000000000
31.5700000000000
EQUILIBRIUM STAGE DISTILLATION SIMULATION
COMPONENTS:
1:H20 2:SBA 3:DiiB
NUMBER OF STAGES 16 DISTIl,LATE RATE 77.120 RE FLUX RATIO 4.200 TOTA1 PRESSURE 760.000
STREAM FLOW RATE TeC) COMPONENT FLOWS
BOTTOMS 17.19 99.6 0.0000 17.1896 0.0004 DISTILLATE 77.12 73.7 31.5800 13.9704 31.5696
STAGE T(C) LIQUID FLOW COMPONENT FLOWS
1 99.60 17.19 0.000 17.190 0.000 2 99.60 418.21 0.002 418.175 0.037 3 99.58 418.21 0.009 418.070 0.134 4 99.51 418.21 0.037 417.687 0.490 5 99.25 418.21 0.147 416.288 1.779 6 98.37 418.21 0.580 411.242 6.392 7 95.60 418.21 2.269 393.887 22.059 8 89.25 418.21 8.611 342.718 66.886 9 84.85 373.12 4.443 222.825 145.856
10 84.26 373.12 1.900 153.336 217.888 11 84.24 373.12 1. 488 138.482 233.154 12 83.95 373.12 1. 759 138.637 232.728 13 83.94 323.90 1. 633 123.504 198.767 14 83.89 323.90 1.674 123.294 198.936 15 82.91 323.90 2.525 122.902 198.477 16 73.66 323.90 17.144 115.989 190.771
FLOW CONFIGURATION
I FL FV SL SV FKV FEEDSTREAMS
1 17.2 401.0 0.0 0.0 0.0 0.0 0.0 0.0 2 418.2 401.0 0.0 0.0 0.0 0.0 0.0 0.0 3 418.2 401. 0 0.0 0.0 0.0 0.0 0.0 0.0 4 418.2 401.0 0.0 0.0 0.0 0.0 0.0 0.0 5 418.2 401.0 0.0 0.0 0.0 0.0 0.0 0.0 6 418.2 401.0 0.0 0.0 0.0 0.0 0.0 0.0 7 418.2 401.0 0.0 0.0 0.0 0.0 0.0 0.0 8 418.2 401.0 0.0 0.0 0.0 26.9 18.2 0.0 9 373.1 401.0 0.0 0.0 0.0 0.0 0.0 0.0
10 373.1 401.0 0.0 0.0 0.0 0.0 0.0 0.0 11 373.1 401.0 0.0 0.0 0.0 0.0 0.0 0.0 12 373.1 401.0 0.0 0.0 0.0 4.7 13.0 31.6 13 323.9 401.0 0.0 0.0 0.0 0.0 0.0 0.0 14 323.9 401.0 0.0 0.0 0.0 0.0 0.0 0.0 15 323.9 401.0 0.0 0.0 0.0 0.0 0.0 0.0 16 323.9 77.1 0.0 0.0 0.0 0.0 0.0 0.0
85
* * * * * * * * * * * * * * * * * * * * K-FACTOR PROFILE IN COLUMN 72.'3
1 4.141 1.000 3.805 ') 4.141 1.000 3.804 .:...
3 4.140 0.999 3.801 4 4.136 0.996 3.789 5 4.123 0.987 3.746 6 4.079 0.955 3.599 ,.,
3.958 0.862 3.162 , 8 3.798 0.681 2.274 9 6.035 0.644 1.390
10 13.911 0.846 0.996 11 17.937 0.938 0.929 12 17.568 0.923 0.921 13 16.444 0.898 0.937 14 16.457 0.897 0.934 15 15.588 0.854 0.905 16 7.737 0.506 0.695
NUMBER OF STAGES 12
NUMBER OF FEEDS 1
THE STAGE AT WHICH FEED 1 IS INTRODUCED 6
THE VAPOR FRACTION OF FEED 1 0.000000000000000
COMPONENT FLOW RATES IN FEED 1 28.250000000000000 0 . 880000000000000
CONDENSER ( YI N)? Y
THE DISTILLATE RATE 3.840000000000000
THE RE FLUX RATIO 4.200000000000000
NUMBER OF LIQUID SIDE STREAMS o
NUMBER OF VAPOR SIDE STREAMS o
THE PRESSURE 760.000000000000000
ESTIMATE THE TOP AND BOTTOM STAGE TEMPERATURES I N DEGREES CELSIUS
73.660000000000000 99.600000000000000
THE MAXIMUM CHANGE IN TEMPERATURE BETWEEN ITERATIONS ( DEGREES CELSIUS ) - OFTEN 10
2.000000000000000
TH E MAXIMUM FRACTIONAL CHANGE IN FLOW RATES BETWEEN ITERATIONS - OFTEN 0.5
0.100000000000000
1.4700000000000
EQUILIBHIUM STAGE DISTILLATION SIMULATION
COMPONENTS:
1:H20 2:SBA 3:DiiB
NUMBEH OF STAGES 12 DISTILLATF. HATE 3.840 HEFLlJX HATIO 4.200 TOTAL PHESSUHE 760.000
STHEAM FLOW HATE T (G) COMPONENT FLOWS
BOTTOMS 26.76 101.5 26.7579 0.0021 0.0000 DISTILLATE 3.84 77.2 1.4921 0.8779 1.4700
STAGE T(C ) LIQUID FLOW COMPONENT FLOWS
1 101.46 26.76 26.758 0.002 0.000 2 101.43 46.73 46.717 0.011 0.000 3 101.35 46.73 46.697 0.031 0.000 4 101. J5 46.73 46.646 0.082 0.000 5 100.54 46.73 46.514 0.208 0.006 6 97.10 46.73 46.117 0.508 0.103 M 96.96 16.13 15.901 0.191 0.036 , 8 96.22 16.13 15.813 0.277 0.038 9 93.24 16.13 15.361 0.720 0.047
10 87.84 16.13 13.557 2.475 0.096 11 84.42 16.13 10.326 5.467 0.335 12 77.19 16.13 8.332 6.187 1.609
FLOW CONFIGUHATION
I FL FV SL SV FKV FEEDSTHEAMS
1 26.8 20.0 0.0 0.0 0.0 0.0 0.0 0.0 2 46.7 20.0 O. 0 0.0 0.0 0.0 0.0 0.0 3 46.7 20.0 0.0 0.0 0.0 0.0 0.0 0.0 4 46.7 20.0 0.0 0.0 0.0 0.0 0.0 0.0 5 46.7 20.0 0.0 0.0 0.0 0.0 0.0 0.0 6 46.7 20.0 0.0 0.0 0.0 28.3 0.9 1.5 7 16.1 20.0 0.0 0.0 0.0 0.0 0.0 0.0 8 16.1 20.0 0.0 0.0 0.0 0.0 0.0 0.0 9 16.1 20.0 0.0 0.0 0.0 0.0 0.0 0.0
la 16.1 20.0 0.0 0.0 0.0 0.0 0.0 0.0 11 16. 1 20.0 0.0 0.0 0.0 0.0 0.0 0.0 12 16.1 3.8 0.0 0.0 0.0 0.0 0.0 0.0
* * * * * * * * * * * * * * * * * * * * K-FACTOR PROFILE IN COLUMN T'2...J
1 1.000 5.986 40.837 2 0.999 5.976 40.778 3 0.997 5.951 40.627 4 0.991 5.888 40.238 5 0.974 5.705 39.128 6 0.883 4.923 34.341 7 0.879 4.880 34.069 8 0.861 4.658 32.659 9 0.791 3.761 26.850
10 0.704 2.071 15.221 11 0.768 1.044 7.431 12 0.752 0.596 3.837
A-5 Program listing and output for
MEK convers ion reactor
----------------------------------------Listing of MEK . PAS, page 1 at U1:~3pm U~/l1/~4
1: proeram MEK_conversion; 2: 3: const 4: MS 74.123;
= ?2.10?; = 2.016;
= 8.3144; = 0.32;
5: MM 6: MH 7: R 8: E 9:
10: Var 1 1 : 12: 13 : 14: 15: 16: 17: 18 : 19: 20: 21 :
T,Dk,Dp,Ks,Km,Kh,Kp,KO,W,dW, Xs,Xm,Xh, F,FO,Fs,Fm,Fh,FsO, P,PO,Ps,Pm,Ph, Ue,Conv,Rho,Ra,TubeLength, TubeDia,Tau,G,H,S,Gr i,j,N PrinterEcho Choice
22: Procedure Initialisation; 23: Beein 24: T:=273.16+310; 25: H:=53429.0+3*T; 26 : 5 : = 1 1 . 54 + 6 . 908':\- In ( T) I In ( 1 0) ; 27: G:=H-T*S; 28: Ok:=0.0005; 29: Ks:=5.25e-14*exp( 15.74e3/T); 30: Km:=0.226*exp(0.87e3/T); 31: Kh:=5.25e-14*exp( 15.74e3/T); 32: Kp:=-2790/T+1.510*ln( T) I1n( 10) +1.865; 33: Kp:=exp(ln(10)i~Kp);
34: KO:=1.3/3.6; {mol/kg . s.atm} 35: W:=TubeLength*pi/4*sqr(TubeDia)*718.8; 36: dW:=W/N; 37: End j 38: 39: Procedure Input; 40: Beein
Real; Integer; Boolean; Char;
41: TubeLength:=0.85;FO:=0 . 71;PO:=2.4;Xs:=0.998; 42: Xm:=0.002;N:=1000;TubeDia:=0.10j 43: ClrScr; 44: GotoXYC5,4) ;WriteC 'Tube diameter (m) ') ;ReadCTubeDia); 45: GotoXYC5,5);WriteC'Tube length Cm) ');Read(TubeLength); 46: GotoXY(5,6) ;WriteC 'Initial flow (molis) ') ;Read(FO); 47: GotoXYC5,7) ;Write( 'Initial pressure Catm)') ;ReadCPO); 48: GotoXYC5,8) ;Write( 'Molf'raction SBA ') ;ReadCXs); 49: GotoXYC5,9) ;WriteC 'Molf'raction MEK ') ;ReadCXm); 50: GotoXY(5,10);Write('# steps ');ReadCN); 51: GotoXY(5, 11) ;Write( 'Printout? C YIN)'); 52: Repeat read(kbd,Choice) Until UpcaseCChoice) in ['Y', 'N']; 53: If' Upcase(Choice)='Y' then PrinterEcho:=True else PrinterEcho:=False; 54: P:=PO; 55: F:=FO; 56: FsO:=FO*Xs; 57: Ps:=Xs*P; 58: Pm:=Xm*P; 59: Xh:=O; 60 : Ph:=O; 61: Fs:=Xs*F; 62: Fm:=Xm*F; 63: Fh:=O; 64: If PrinterEcho Then 65: Begin 66: Writeln(lst);
67: 68: 69: 70: 71 : 72: ?3: 74: ?S: 76: 77: 78: 79: 80: 81 : 82: 83: 84: 85: 86: 87: 88: 89: 90: 91 : 92: 93: 94: 9S: 96: 97: 98: 99:
100: 101 : 102: 103: 104: 10S: 106: 107: 108: 109: '10 : 111 : 1 12 : 1 13: , 14 : 11S: 1 16 : 117: 118 : 119: 120: 121 : 122: 123: 124: 12S: 126: 127: 128 :
- -- _._---_. ------------
Listing of MEK.PAS, page 2 at 01:33pm 07/11/84
Writeln( lst, ' Writeln( lst, ' Writeln( lst, , Writeln( lst,' Writeln( lst, , Writeln( lst, '
End; End;
Procedure Output; Begin
ClrScr;
Tube diameter Tube length Initial flow Initial pressure Molfraction SBA # steps
" TubeOia:7:4,' m'); ',TubeLength:7:4,' m'); ',FO:7:4,' mol/s'); ',PO:7:4,' atm'); ',Xs:7:4); , ,N:7);
GotoXY(S,4);Write(j,' pressure = ',P:7:4,' atm'); GotoXY(5,S);WriteC' convers ion ',CFO-Fs)/FO:7:4); GotoXY(S,6) ;WriteC' rho = ',Rho:7:4,' kg/m3'); GotoXY(S,7) ;Write(' flow ' ,F:7:4,' mol/s'); GotoXY(S,8) ;WriteC 'reaction enthalpy = ',Gr/1000:7:4,' kW'); GotoXY(S,9) ;WriteC' x(SBA) = ',Xs:7:4); GotoXY(S,10);Write(' x(MEK) = ',Xm:7:4); GotoXY(5,11);Write(' xCH2) = ',Xh:7:4); GotoXY(S,20) ;WriteC' Press key to continue '); If Not PrinterEcho then repeat until keypressed; GotoXY(S,20) ;WriteC' '); If PrinterEcho Then Begin Writeln( lst); Writeln( lst,' Write(lst,' Writeln( lst,' Write(lst,' Writeln( lst, ' WriteCIst,' WritelnC lst, ' WritelnC lst, , End;
j*(N div S):4,': pressure = ',P:7:4, conversion = ',Fh/FsO:7:4);
x(SBA) = ' ,Xs:7:4); rho = ' ,Rho:7:4,' kg/m3');
x(MEK) = ',Xm:7:4); flow = ' ,F:7:4,' mol/s');
x(H2) = ',Xh:7:4); reaction enthalpy = ' ,Gr/1000:7:4,' kW');
End;
Procedure Kinetics; Begin
Ra:=KO*(Ps-(Pm*Ph/Kp))/( 1+Ks*Pm+Km*Pm*Kh*Ph); Conv:=dW*Ra;
End;
Procedure GasOensitYi {kg/m3} Begin
P: =P-:l-1 0 132S . 0; Rho:=P/R/T*CXs*Ms+Xm*Mm+Xh*Mh)/1000; P:=P/10132S.0;
End;
Procedure GasVelocity; {superficial} Begin
Ug:=F/(pi/4i~sqrCTubeOia))*(Xs*Ms+Xm*Mm+Xh*Mh)/Rho/1000;
End;
Procedure PressureOrop; Begin
GasOensity; GasVelocity; Op:=1.7S*( 1-E)/(E*E*E)*Rho*sqr(Ug)*TubeLength/N/Ok/10132S.0;
End;
129: Begin 130: Input; 131: Initialisation; 132: Clr5cr;
133: 134: 135: 136: 13?: 138: 139: 140: 141 : 142: 143: 144: 145: 146: 14?: 148: 149: 150: 151 : 152:
Listin2 of MEK.PAS, pa2e 3 at 01:33pm 0?/11/84
For j:=1 to 5 do Begin
For i:=1 to (N div 5) do Begin
GasDensity; GasVelocity; Tau:=TubeLength/N/Ug; Kinetics; Fs:=Fs-Conv;Fm:=Fm+Conv;Fh:=Fh+Conv; F:=Fs+Fh+Fm; Gr:=G*Fh; Xs:=Fs/F;Xm:=Fm/F;Xh:=Fh/F; Xs:=Fs/F;Xm:=Fm/F;Xm:=Fm/F; PressureDrop; P:=P-dP; Ps:=P*Xs;Ph:=P*Xh;Pm:=P*Xm;
End; Output;
End; End.
Tube diameter O. 1000 m Tube length 0.8500 m Initial flow 0.?100 mol/s Initial pressure 2.4000 atm Molfraction S8A 0.9980 # steps 1000
200: pressure 2.2419 atm conversion 0.5613 x(5BA) = 0.28o?
rho = 2.226? kg/m3 x(MEK) = 0.3603 flow 1 . 1o?? mol/s x( H2) 0.3590
reaction enthalpy = 14.83?0 kW
400: pressure 2.0294 atm conversion = 0 . ?456 x(SBA) = O. 1456
rho 1 .8034 kg/m3 x( MEK) = 0 . 42?8 flow = 1.2383 mol/s x( H2) 0.4266
reaction enthalpy = 19.?086 kW
600: pressure = 1 . ??26 atm conversion = 0 . 8310 x( S8A) 0 . 0922
rho = 1 . 5021 kg/m3 x( MEK) = 0.4544 flow = 1 .2988 mol/s x( H2) = 0 . 4533
reaction enthalpy 21.9661 kW
800 : pressure = 1 . 4602 atm conversion = 0 . 8?62 x( S8A) 0.0659
rho 1 . 2080 kg/m3 x( MEK) 0.46?6 flow = 1 .3308 mol/s x( H2) 0.4665
reaction enthalpy = 23.1615 kW
1000: pressure = 1 .0505 atm conversion = 0.9020 x(S8A) = 0.0515
rho 0.8583 kg/m3 x(MEK) 0.4?48 flow 1 . 3491 mol/s x( H2) = 0 . 4?3?
reaction enthalpy = 23 . 8442 kW
34
Tube diameter O. '000 m Tube length 0.8500 m Initial flow 1 .4200 mol/s Initial pressure 4.4000 atm Molfraction BBA 0 . 9980 # steps 1000
200: pressure = 4.0596 atm conversion = 0.5209 xC BBA) = 0.3146
rho = 4.1394 Iq~/m3 xC MEK) = 0.3434 flow = 2. 1582 mol/s xC H2) = 0.3420
reaction enthalpy = 27.5389 kW
400: pressure = 3 . 5986 atm conversion = 0.6965 xC BBA) = 0.1787
rho = 3.2905 kg/m3 xC MEK) = 0.4113 flow = 2.4071 mol/s xC H2) = 0.4101
reaction enthalpy 36.8259 kW
600: pressure 3.0254 atm conversion = 0 . 7816 xC BBA) = 0.1224
rho 2 . 6354 kg/m3 xC MEK) = 0.4393 flow 2.5277 mol/s xC H2) 0 . 4382
reaction enthalpy 41 .3232 kW
800: pressure 2.2844 atm conversien = 0.8289 xC BBA) 0.0935
rho = 1 .9401 kg/m3 x C MEK) 0 . 4538 flow 2.5947 mol/s x( H2) = 0.4527
reaction enthalpy = 43 . 8227 kW
1000: pressure = 1 .0998 atm conversion = 0.8556 x( BBA) = 0.07?8
rhe = 0.9265 kg/m3 x( MEK) = 0.4617 flow = 2.6325 mol/s x( H2) = 0.4606
reaction enthalpy = 45.2333 kW
A-6 Stream data MEK purification column T43
VERSION 0484
SM PROCESS INPUT LISTING - PAGe 1
GENERAL DATA TIlLE USER=]\ AND R' , PROBLEM=MEK PURIF. ,PROJECT=FVO, DATE=E'ED87 DIMENSION SI,TEMP=C,PRESS=BAR P RINl WTO PTION
COMPOriENl DATA LIBID 1,MEK/2,SBUOH/3,~ATER
THERHODYNAMIC DATA TYPE SYSTEH=SRK
SlREAM DATA PROPERTY STRM=FD,PRESS=1.0,PHASE=L,* COMP(M)=1,89.46/2,10.04/3,O.SO,NOCHECK,RATE(M)=72.07
UNIl OPERATIONS COLUMN UID=ACOL,KPRINT
PARAMETER TRAY=20,FAST=5,SURE=30,DKDX FEED FD,7 PRODUCT OVHD=ATOP,65,BTMS=ABOT,7.8 HEATER 1,20,4/2,1,-3 CONDENSER TYPE=3,PRESSURE=1.0 PSPEC TOP=1.01,DP=O.01 SPEC STRM=ATOP,COMP=1,1,FRACIION(Y) =0.99 VARIA3LE HEAT=1 PRIrJT TRAY=20 PLOT PROFIL~,XCOMP=1/2/3,YCOHP=1/2/3 ESTIMATE TTEMP=7S.0,CTEMP=7S.0,BTEMP=100,RTEMP=100
VERSION 0404 P~GE 8 SIMULATION SCIENCES, INC.
PROJECT FVO
SM PROCESS
UNIT 1 - ACOl SOlUTION
A, AND H PROBLEM MSK PURIF. FE.JJ7
I SUHMARY FOR COLUMN
1 TOT Al Nut1BER OF ITERATIONS FAST METHOD SURE !1ETHOD
2 cOlurm SUMMARY
UNIT 1 - ACOL"
o 7
TRAY TEMP PRESSURE NET FLO~ RATES, KG MOLS/Ha HEAT (COOL) ER LIQUID VAPOR DUTIES
DEG C BAR PHASE(L) PHASE(V} FEED PRODUCT MM KJ /HH
1. 2 3 4 5 6 7 8 9
10 11 12 13 14 15 1.6 11 18 19 20
18.3 79.5 80.1 80.6 81.1 81.7 82.3 82.7 83.0 83.3 83.6 83.9 84.2 84.5 84.8 85.1 85.4 85.8 66.5 81.7
1.00 1.01 1.02 1.03 1.04 1.05 1.06 1.07 1.08 1.09 1.10 1.11 1.12 1.13 1.14 1.15 1.16 1.1.1 1.18 1.19
59.7 59.7 59.5 59.3 59.1 58.9 ) /
131.7 . 1.31.8 131.9 132.1 132.2 132.3 132.4 132.5 132.5 132.5 132.5 132~3 131.8
3 FEED Arm PRCDUCT STREAMS * FEED STREAMS:
FD TO TRAY 7 IS LIQUID * PRODUCT STREAMS:
92.1 92.1 92.0 91.8 91.6 91.3 92.1 92.2 92.3 92.5 92.6 92.7 92.8 92.8 92.9 92.9 92.9 92.7 92.2
ABOT IS lIQUID STREAM FROM TRAY 20 ATOP IS lIQUID STREAM FROM TRAY 1
32.5L
72.1L
39.6L
MASS RATES KG MOlS/HR
0.72070E+02
0.39603E+02 0.32467E+02
O.OOOOOE+OO
-3.0000
3.0331
HEAT RAIES MM KJ /Hf<
0.91590E+00
o .58 0 7 9E +00 0.37439E+00
OVERALL HASS BAlANCE, (FEEDS - PRCDS) OVERALL HEAT BALANCE" (HIN - HOUT) -0.62237E-02
4 SPECIFICATION VAlUES
PARA!1ETER TRAY COMP. SPECIFICATION TYPE NO NO TYPE
STRM ATOP 1 1 \J.F •
:17
SPECIFIED VALUE
0.9900E+00
CALCULATED VALUE
0.9896E+00
VERSION 0404 SM SIMULATION SCIENCES. INC. PROCESS PA GI:: 9 PROJECT FVO UNIT 1 - ACOL A. ANC R. PROBLEN MeK PURIF. SOLUTION FEB67
IIA TRAY COMPOSITIONS
TRAY -------- 1 -------- -------- 2 --------COMPONENT X Y X Y 1 MEK 0.98l5E+OO O.9615E+OO o .9626E+00 o .9611E+00 2 SBUOH 0.7347E-02 O.7347E-02 0.l466E-Ol 0.7344E-02 3 YATER 0.11l0E-Ol o .1110E-Ol 0.2763E-02 O.lllOE-Ol
KG t10LS/HR O.5967E+02 0.3247E+02 o .5966E+02 O.9213E+02
TRAY -------- 3 ------- -------- 4 -------COMPOHENT X Y X y . 1 MEK 0.9747E+00 O.9617E+OO 0.9636E+00 O. 9766E +00 2 SBUOB 0.2367E-OI O.1209E-Ol O.3S25E-Ol O.1603E-01 3 \.lATER o .1412E-02 O.5699E-Q2 o .l19SE-02 O.4629E-02
KG HOLS/HR O.59S0E+02 0.92l2E+02 0.5930E+02 0.9l97E+02
TRAY -------- 5 ------- -------- 6 --------COHPONENT X Y X Y 1 MEK 0.9497E+00 0.9696E+OO O.9333E+00 0.9606E+00 2 SBUOH O.4909E-Ol O.2S37E-Ol 0.6SS8E-Ol O.3429E-Ol 3 WATER o .ll60E-02 0.4697E-02 o .llSSE-02 O. 4663E -02
KG HOLS/HR O.S909E+02 0.9l77E+02 o .5666E+02 O. 9lS5E +02
TRAY -------- 7 -------- -------- e -------COMPONENT X Y X Y 1 HEK 0.9l4IE+OO O.9504E+OO 0.9148E+OO O.9531E+OO 2 SBUOH O.647lE-Ol 0.4469E-Ol 0.6460E-01 O.4Sl6E-01 3 WATER 0.ll53E-02 O.4669E-02 o .405lE-03 o .1649E-02
KG MOLS/BR 0.l317E+03 0.9133E+02 0.l3l6E+03 O. 9207E +02
TRAY -------- 9 -------- -------- 10 --------COHPONENT X Y X Y. 1 MEK 0.9l50E+OO 0.9540E+00 0.9l50E+00 0.9542E+00 2 SBUOH o .6466E-Ol 0.4534E-Ol 0.6493E-Ol 0.4S49E-01 3 WATER 0.l424E-03 0.579lE-03 o .SOO7E-04 0.2034E-03
KG ~10LS/HR 0.13l9E+03 O.922lE+02 O.l32lE+03 O.9233E+02
TRAY -------- 11 -------- -------- 12 --------COMPONENT X Y X Y 1 MEK 0.9l50E+OO O.9542E+OO O.9l49E+OO o .954lE+00 2 SBUOH o .6S00E-Ol O.4563E-Ol o .6Sl0E-Ol O.4579E-01 3 WATER O.l763E-04 O.7l53E-04 o .621SE-05 O.2517E-04
KG HOLS/HR o .1322E+03 O.9246E+02 o .1323E+03 0.9257E+02
VERSION 0484 SH SIHULATION SCIENCES, INC. PROCESS PAGE 10 PROJECT FVO UNIT 1 - ACOl A. AND R PROBLEM MSK PURIF. SOlUIION fEB87
TRAY -------- 13 -------- -------- 14 --------COMPONENT X Y X Y 1 HEK 0.9147E+00 0.9539E+00 0.9144E+00 0.9536E+00 2 SDUOH o .8527E-01 0.4597E-01 0.8559E-01 0.4624E-Ol 3 WATER 0.2193E-05 O.8870E-05 0.7747E-06 o. 3129E -05
KG 110lS/HR 0.1324E+03 0.9267E+02 0.1325E+03 0.9277E+02
TRAY -------- 15 -------- -------- 16 --------CDt1PONENT X Y. X Y 1 HEK 0.9137E+00 0.9531E+00 0.9121E+OO 0.9520E+00 2 SBUOH 0.8628E-Ol O.4673E-Ol 0.8791E-01 O.477SE-Ol 3 WATER 0.2739E-06 O.110SE-05 o .9682E-07 0.3900E-06
KG ~1OlS/HR 0.1325E+03 0.9285E+02 0.1325E+03 O. 9291E+ 02
TRAY -------- 17 ------- -_._----- 18 --------COHPONEN! X Y X Y 1 HEK 0.9082E+00 0.9497E+00 O.8986E+00 0.9440E+00 2 SBUOH 0.9185E-01 o .5009E-01 o .1014E+00 0.5569E-Ol 3 WATER 0.3416E-07 0.1375E-06 0.1196E-07 0.4816E-07
KG HOlS/HR 0.1325E+03 0.9294E+02 0.1323E+03 0.9288E+02
TRAY -------- 19 -------- -------- 20 --------COt1PONENT X Y X Y 1 MEK 0.8758E+00 0.9307E+00 0.8234E+00 O. 8989E +00 2 SBUOH o .1242E+00 0.6924E-01 o .1766E+00 o .1017E +00 3 ~ATER o .40 84E-08 o .1651E-07 o .1293E-08 0.5285E-08
KG 110LS/HR 0.1318E+03 0.9269E+02 0.3960E+02 0.9216E+02
VERS ION 0484 SIMULATION SCIENCES, INC. PROJECT FVO PROBLEH HEK PURIF.
SH PROCESS
SOLUIION
STREAH SUMMARY
STREAH ID. NAHE PHASE
FROM UNIT/TRAY Ta UNIT/TRA Y
FROM STREAt1
KG MOlS/HR I~1PERATURE, DEG C PRESSURE, BAR H, MM KJ /HR
H KJ /KG MOlE KJ /KG
MOlE FRACT LIQUIO
11 KGS/HR MOLECULAR ~EIGHT STO LIQ IB/HR
UOP K
OEG API SP GR KGS/l-!3
REOUCEO TEMP REOUCED PRESS ACENIRIC FACTOR *:::VAPOR**
\JEIGHT M3/HR
M3/HR
M KGS/HR t10lECULAP. STO LIQ STO M ACIUAL H H3/HR
KGS/H 113 Z CP,KJ /KG MOL C
::::::~IQUID**
t1 KGS/BR MOLECULAR \JEIGHT STO LIQ M3/HR ACTUAL GPl1
Z
H3/HR KGS/H3
CP,KJ /KG Mal C
FO
LIQUID Ol 0 11 7
72.070 60.413 1.060 0.916
12.706 176.409 1.00000
5.192 72.040
6.423 43.193 0.6100
806.3293 10.632
0.659 0.025 0.350
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 O.OOOOE+OO
5.192 72.040
6.423 37.3663
6.487 611.761 0.00425
1.7423E+02
/00
ABOT
LIQUID 1/ 20 Ol 0
39.603 87.652 1.190 0.581
14.665 202.361 1.00000
2.670 72.464
3.551 43.231 0.6096
806.1536 10.649
0.674 0.029 0.369
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 O.OOOOE+OO
2.670 72.464 3.551
20.7729 4.716
606.261 0.00473
1.7913E+02
ATOP
:IQUID 1/ 1 0/ 0
32.467 76.313 1.000 0.374
11.532 1G1.231 1.00000
2.322 71.522
2.672 'D.145 0.6102
808.5463 10.611
0.655 0.023 0.326
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 O.OOOOE+OO
2.322 71.522
2.872 16.7752
3.610 609.462 0.00402
1.6861E+02
PAGE 17 Al . At~D Fr FEB67
----- ._.--_.-- - --------------
A-7 Stream data MEK purification column T5l
VERSION 0484
SM PROCESS INPUT LISTING - PAGE 1
GENERAL DATA TITLE USER=A AND R ,PROBLEM=MEK PURIF.,PROJECT=FVO,DATE=FED87 OlMEN SION S1,TEMP=C,PRESS=BAR PRINT WTOPTION
COMPONENT DATA L1BIO 1,MEK/2,SBUOH
THERMOOYNAMIC DATA TYPE SYSTEM=SRK
STREAM DATA PROPERTY STRM=FD,PRESS=1.0,PHASE=L,~ COMP(M)=1,82.338/2,~7.662,NOCHECK,RATE(M)=39.6032 UNIT OPERATIONS
COLUMN UID=BCOL,KPR1NT PARAMETER TRAY=25,FAST=5,SURE=30,DKDX FEED FD,7 _ PRO DUCT OVHD=ATOP, 32.6085, BTMS =ABOT , 6 .99" 7 HEATER 1,25,3/2,1,-4.0 CONDENSER TYPE=3,PRESSURE=1.0 PSPEC TOP=1.01,DP=0.Ol SPEC STRM=ABOT,COMP=2,2,FRACTION(W) =0.99 VARIABLE HEAT=l PRINT TRAY=25 PLOT PROFILC,XCOMP=1/2,YCOMP=1/2 ESIINATE ITEMP=75.0,CTEMP=75.0,BTEMP=100,RTEMP=100
lOl
VERSION 0484 SIMULATION SCIENCES, INC. PROJECT FVO PROBLEM MEK PURIF.
I SUMMARY FOR COLUMN
1 TOTAL NUMBER OF ITERATIONS FAST HETHOO SURE METHOO
2 COLUMN SUMMARY
SM PROCESS
UNIT 1 - BCOL SOLUTION
UNIT 1 - BCOL,
PAGE 8 A AND R FEB87
o 4
TRAY TEMP PRESSURE NET FLOU RATES, KG MOLS/HR HEAT(COOL)ER LIQUIO VAPOR DUTIES
DEG C BAR PHASE(L) PHASE(V) FEED pnODUCT MM KJ IHR
1 79.4 2 79.8 3 80.3 4 80.8 5 81.4 6 82.1 7 82.9 8 83.2 9 83.6
10 84.1 11 84.6 12 . 85.4 13 86.4 14 87.9 15 89.9 16 92.3 17 95.0 18 97.5 19 99.7 20 101.4 21 102.6 22 103.5 23 104.2 24 104.7 25 105.1
1.00 1.01 1.02 1.03 1.04 1.05 1.06 1.07 1.08 1.09 1.10 1.11 1.12 1.13 1.14 1.15 1.16 1.17 1.18 1.19 1.20 1.21 1.22 1.23 1.24
90.6 90.4 90.2 89.9 89.5 89.0
128.9 128.9 128.8 128.6 128.2 127.4 126.2 124.4 122.0 119.~
116.9 114.8 113.3 112.3 111.7 111.3 111.1 111.0
3 FE EO AND PRODUCT STREAMS * FEED STREAMS:
FD TO TRAY 7 IS LIQUID * PRODUCT STREAMS:
123.4 ' 123.2 123.0 122.7 122.3 121.7 122.1 122.1 122.0 121.8 121.3 120.6 119.4 117.5 115.2 112.5 110.0 108.0 106.5 105.5 104.8 104.5 104.3 104.2
ABOT IS LIQUIO STREAM FROM TRAY 25 ATOP IS LIQUIO STREAM FROM IRAY 1
OVERALL MASS BALANCE, (FEEOS - PROOS) OVERALL HEAT BALANCE, (HIN - HOUI)
10'2..
32.8L
39.6L
6.8L
MASS RAIES KG MOLS/HR
0.39603E+02
0.68389E+01 0.32764E+02
O.OOOOOE+OO
-4.0000
4.0153
HEAT RATES MM KJ IHR
0.54223E+00
o .17689E+00 o .3B061E+OO
0.15914E-04
VERSION 0484 SM SIMULATION SCIENCES, INC. PROCESS PAGE 10 PROJECT FVO UNIT 1 - BCOL A. AND R PROBLEH MEK PURIF. SOLUTION FEB87
IIA TRAY COMPOSITIONS
TRAY -------- 1 -------- -------- 2 --------COHPONENT X Y X Y 1 MEK 0.9931E+OO 0.9931E+00 0.9864E+00 0.9931E+00 2 SBUOH 0.6903E-02 0.6903E-02 0.1365E-01 0.6903E-02
KG MOLS/HR 0.9059E+02 0.3276E+02 o .9045E+02 0.1234E+03
TRAY -------- 3 ------- -------- 4 --------COMPONENT X Y X Y 1 MEK 0.9767E+OO 0.9881E+00 o .9633E+OO O.9811E+OO 2 SBUOH O.2326E-01 O.1185E-01 0.3673E-01 O.1890E-01
KG MOLS/HR O.9023E+02 O.1232E+03 0.8993E+02 0.1230E+03
TRAY -------- 5 ------- -------- 6 --------COHPONENT X Y X Y 1 MEK 0.9448E+00 0.9712E+00 0.9202E+OO O. 9578E +00 2 SBUOH 0.5518E-01 o .2876E-01 0.1915E-01 0.4225E-01
KG HOLS/HH 0.8951E+02 O.1221E+03 0.8896E+02 0.1223E+03
TRAY -------- 7 -------- -------- 8 --------COHPONENT X Y X Y 1 MEK O.8887E+OO O.9399E+00 0.8856E+OO O.9379E+OO 2 SBUOH 0.1113E+00 0.6015E-01 o .1144E+OO 0.6207E-01
KG MOLS/HH 0.1289E+03 0.12l1E+03 0.l289E+03 0.l22lE+03
TRAY -------- 9 ------- -------- 10 --------COt1PONENT X Y X Y 1 MEK 0.8802E+00 0.93Q7E+00 0.8101E+00 0.9290E+OO 2 SBUOH 0.1l98E+OO 0.6535E-01 0.l293E+00 o .1l05E-01
KG MOLS/HR 0.l288E+03 0.l22lE+03 o .l286E+03 o .1220E +03
TRAY -------- 11 ------- -------- 12 --------COMPONENT X Y X Y 1 MEK 0.8543E+00 0.9l90E+00 0.8261E+00 0.9019E+OO 2 SBUOH 0.1457E+OO O.8097E-Ol o .l133E+00 0.98l0E-Ol
KG MOLS/HR 0.l282E+03 o .l2l8E+03 0.l27QE+03 0.12l3E+03
TRAY -------- 13 ------- -------- 14 --------COMPONENT X Y X Y 1 MEK O.7818E+00 0.8130E+00 0.7l36E+OO O.8260E+OO 2 SBUOH 0.2182E+OO o .1210E+00 o .2864E+OO O.1140E+OO
KG MOLS/HR 0.1262E+03 o .1206E+03 o .l2Q4E+03 o .1194E+03
103
VERSION 0484 SM SIMULATION SCIENCES. INC. PROCESS PAGE 11 PROJECT FVO UNIT 1 - BCOL A AND R PROBLEH HEK PURIF. SOLUTION FEB87
TRAY -------- 1S -------- -------- 16 --------COMPONEN:: X Y X Y 1 MEK 0.619SE+00 0.7S46E+00 o .SOSOE+OO O.6SS7E+OO 2 SBUOH 0.380SE+OO O.24S4E+OO O.49S0E+00 O.3443E+OO
KG HOLS/HR 0.1220E+03 O.117SE+03 O.1194E+03 o .11S2E+03
TRAY -------- 17 ------- -------- 18 --------COHPONENT X Y X Y 1 MEK O.3843E+00 O.S3S1E+OO O.2142E+OO O.407SE+00 2 SBUOH o .61S7E+OO 0.4649E+00 O.72SBE+00 O.S924E+OO
KG .HOLS/HR 0.1169E+03 0.112SE+03 o .1lC~8E+03 O.1100E+03
TRAY -------- 19 -------- -------- 20 --------CQt'1PONENT X Y X Y 1 MEK 0.18SSE+00 0.2909E+OO 0.1206E+00 0.1961E+00 2 SBUOH o .814SE+00 0.7091E+00 0.8194E+OO o .8032E+00
KG ~mLS/HR 0.1133E+03 0.1080E+03 o .1123E+03 o .106SE+03 ·
TRAY -------- 21 -------- -------- 22 --------COHPONENT X Y X Y 1 MEK 0.7622E-01 O.1277E+00 O.4125E-01 O. 80S2E -01 2 5BUOH 0.9238E+00 0.8723E+OO O.9S27E+OO O. 9195E+ 00
KG l·10LS/HR 0.1117E+03 o .105SE+03 o .1113E+03 0.1048E+03
TRAY -------- 23 -------- -------- 24 --------COMPONENT X Y X Y 1 MEK o .288SE-01 O.4967E-01 0.1731E-01 0.3007E-01 2 5BUOH 0.9711E+00 0.9S03E+OO 0.9826E+00 0.9699E+00
KG MOL5/HR 0.1111E+03 O.104SE+03 0.1110E+03 O.1043E+03
TRAY -------- 25 --------COMPONENT X Y 1 MEK o .l027E-Ol o .1783E-Ol 2 5BUOH 0.9897E+OO O.9822E+00
KG MOLS/HR O.6839E+01 O.10Q2E+03
VERSION 0484 SIMULATION SCIENCES, INC. PROJECT FVO
SM PROCESS
PROBLEM MEK PURIF.
STREAH 10. NAtiE PHASE
FROM UNIT/TRA Y TO UNIT/TRAY
FROM STREAM
KG l-tOLS/HR rEMPERATURE. DEG C PRESSURE, BAR H, MM KJ /HR
M KJ /KG MOLE KJ /KG
MOLE FRACT LIQUID
M KGS/HR MOLECULAR WEIGHI STO LIQ M3/HR
UOP K
OEG API SP GR KGS/M3
REDUCEO IEMP REDUCEO PRESS ACENIRIC FACTOR **VAPOR**
M KGS/HR MOLECULAR WEIGHT SlO L1Q H3/HR SlO M M3/HR ACIUAL M M3/HR
KGS/M M3 Z CP,KJ /KG MOL C
*::rLIQUIO~
M KGS/HR MOLECULAR WEIGHT SlO LIQ M3/HR ACIUAL GPM
Z
M3/HR KGS/M3
CP,KJ /KG MOL C
SOLUI10N
FO
STREAM SUMMARY
ABOT
LIQUIO 0/ 0 1/ 7
39.603 82.184
1.060 0.5'12
13.692 188.94'1 1.00000
2.870 72.464
3.551 "3.231 0.8098
808.1538 10.6'19
0.663 0.025 0.369
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 O.OOOOE+OO
2.870 72."64
3.551 20.5947
".678 613.525 0.00424
1.7730E+02
/05
LIQUID 1/ 25 0/ 0
6.839 105.093
1.240 0.177
25.865 3"9.041 1.00000
0.507 74.103 0.627
"3.195 0.8100
.808.3179 10.805 0.706 0.030 0.576
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 O.OOOOE+OO
0.507 74.103 0.627
3.56"0 0.809
626.063 0.00467
2.1822E+02
ATOP
LIQUIO 1/ 1 0/ 0
32.76" 79.378 1.000 0.381
11.617 161.068 1.00000
2.363 72.122
2.92" tB.238 0.8098
800.1187 10.616
0.658 0.024 0.326
0.000 0.000 0.000 0.000 0.000 0.000
0.00000 O.OOOOE+OO
2.363 72.122
2.92" 17.0973
3.883 608.521 0.0040"
1.7023E+02
PAGE 22 A. _ AND R' FEB87
A-8 Utility costs
It is assumed that utility costs exists of cooling water, steam,
electricity and natural gas costs.
Cooling water: Available at a temperature of 20°C.
Steam:
Electricity:
Natural gas:
Maximum allowed temperature 40°C.
Costs: f 0.06/m 3 •
Available at 190°C and 3 bar. Enthalpy change
when condensed to water (100°C and 1 bar) is
2365.4 kJ/kg. Costs: f 40.-/ton.
Costs: f 0.19/kWh.
Lower heating value 37.68 MJ/kg.
Costs: f 14.40/GJ.
Table (1): Cooling water demand
Equipment
no.
H2
T3
H4
H18
H27
H32
H38
H40
H47
H55
H57
H58
Total
Capacity
kW
489
5416
506
1889
",2700
220
302
21
832
1123
59
57
C.w. rate
m3 /hr
21. 064
232.992
21. 758
81. 281
116.172
9.464
12.974
0.891
35.777
48.312
2.556
2.448
585.689
Total annual cooling water costs: f 253,018 3
• /0
Table (2): Steam demand
Equipment
no.
H13
T14
H19
H24
H30
H44
H52
Total
Capacity
kW
1305
27
",1500
",1200
843
1115
Steam rate
ton/hr
1.987
2.952
0.041
2.283
1. 825
1. 282
1. 697
12.067
Total annual steam costs: f 3,475,296
For electricity demand only compressor Cl is taken in acount.
The power demand of 105.3 kW requires annual f 144,050.
For the demand of natural gas two cases have been regarded: in
case 1 the acid-reconcentration unit is fully supported by hot flue
gases (no natural gas demand) while in case 2 this unit is fully
supported by natural gas combustion.
Table (3): Natural gas demand
Equipment
no.
F36
R37
Total case 1
Capacity
kW
225
668
acid-reconc. 9736
Gas flow rate
kg/hr
21. 50
63.82
85.32
930.20
- ---------------------------------
Total case 2 1015.52
Total gas casts: case 1 : f 333,310
Total gas casts: case 2: f 3,967,253
The total annual utility casts for case 1 : f 4,205,674
The total annual utility casts for case 2: f 7,839,617
In the used economic model the utility casts are estimated as
10% of the total production casts of f 51,835,530/yr. This estima
tion can only be justified wh en at least 73% of the required energy
for acid reconcentration is supplied by flue gases.
/o~
H32
V33
C , GAS COMPRESSOfI H 2 CONDENSOR T 3 8UTENE A8SOReER H 4 AC/O COOLER P!I AC/O ~
R_G ..
11 8 RECONCEHTRAT1ON IIUSEl C)o1 GAS-UOUO SE_OR 11 8 RECONCENTRATIOH IlESSEl 11 11 HYDAOLYSIS TANK PlO ACID PUMP .. " IlENTUIII SCR~R
F .... '
~ HUnR Sec. 8UTANOI. STRIPPER PUMP STOIIAGE CAUSTIC SCRU8eER CONlEHSOR . PREHUTER l'O.-l'O. SE""'RATOR . ~
T 23 OISTUATIQH C~ H24 REIIOIl.ER P2!1 REFlUX PUMP 1128 UO. ·l/O. SEP""'ATOR H21 COI;OENSOR P28 PUMP Uil OISTUATIOH CCU~
~~~ER H32 COOLER 1133 AlCOHol. SlORAGE
: -: re : 1"'116 1- :----~::J
• to Sewer
:~~oy~, , H19
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Pl~ PUMP Hl!l HEAT EXCHANGER F3e FUANACE Rl7 MUlT'TUBUlAR REACTOR Hle CONDENSOR
C)o GAS·llOiAO SEPARATQII H ~ CONDENSOR H"' O)NDf.N!iOII P42 ~ Hl O'STIL'. AT'ON COLUMN H~~ RE8ou. ER
D;~ B ~
T23
St •• m
I :::ftl: nc ... '
P", RE FL UX PUMP v~e VESSEL ~7 CONDENSOR P~8 PUMP P"II PUMP P~ REC'lClE PUMP T~l DlSTIllATtON CQl.1"II.4N H!l2 RE80ILER P~l REFl.UX PUMP V~ lIESSEL H!I CONDENSOR
Hydrogen
Ing Wat.,.
PROCESS SCHEME FOR PROOUCnoN OF METHYL ETHYL KETONE
A.H. Amer F: V O. no. 2e8J R.F" .de Ruiter April 1~81
OStreom na. [IJ ~ In -C @Pr ..... In ar
P!III PUMP H~7 COOLER H COOLER
~Yltem~ nol gI .... Is' 80r
Errata
on flow sheet F.V.O. no. 2693:
1 To prevent an inert-build up in T3, a part of stream 7 must be purged.
2 Butene absorber T3 is not a jacketed vessel but a vessel with multitubular internal cooling.
3 Gas-liquid separators Cy7 and Cy39 are not cyclones but horizontal separation tanks.
4 NaOH for acid removal can be exchanged with Ca(OH)z, which is + -
cheaper. CaS04 removal is easier than Na /SO~ removal. 5 Brine-cooling for but ene liquification is only necessary when
normal cooling water is not able to reduce the temperature to 25°C.
6 Liquid-liquid separator V20 is a horizontal vessel.
,//