The production of methyl ethyl ketone from n-butene

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

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.

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

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:

- - - ) ( -CH3-ÇH-C&Hs + 3 H&O ---) OS03 H CH3-ÇH-C&Hs + 2 H&O OS03 H

The absorption of but ene can be carried out in 65 wt-% sulfurie acid at 50-60oC, in 75-80 wt-% acid at 30-50oC 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,

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

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 exploexplo-sion 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*107 _{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

7

,

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

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 SiOz 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

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 C a r e 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

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/m3

-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/m3

-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

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 _ss

=

17

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)

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.) (m3_/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

- - - -- - - -- --

-that a 6T

ln 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 V

tt 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

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 _v

=

0.035 ( (Pl-p )/ P ) _{v v} (9)

where Pv and PI are the densities of respectively vapor and

liquid (kg/m3 _{) . 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 %) 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

(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/m3 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

feed one mole of HZS04 is solved in two moles of water. In the

product stream leaving the hydrolysis tank however, one mole of

HZS04 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

HZS04 and for an acid-water ratio of one to thirteen it

kcal/mole HZS04 [19J. By diluting the acid in the

tank an excess of 6.73 kcal/mol HZS04 (28.20 kJ/mol) is

~

-released. The total heat product ion becomes:

65.5 kmoles/hr HZS04

*

=

=

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 % water, 6.83 % 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

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

=

0.140 , 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

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."/ ..." ) t -SBA 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

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 x

f

=

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

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 HZS04 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.

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 .

.

,

....---__, :w d,f,f:i,n . , . . . . IICI • . . . ~.L. ""' • • •• oovc, ac .. Hw . . . . , _,t •• eo.cI.' •• '''.'

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)

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 _s 1 of 0.460 (77.8 wt-%) and the lower layer has

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/cm3 _{) . The}

dispersed phase is the heavy, water-rich, phase and the viscosity

of water at 900

_e

_{is 0.3147 cP. At 90}0

_e

_{the densities of SBA and HzO}

are respectivily 0.78347 g/cm3 _{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/cm3 _{• The lower layer density has the}

value 0.9388 g/cm3 _{• 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

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

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.

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)

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

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' 6H

To and 6S

To were found as function of the temperature (T in K): log K -2790 + 1. 510

*

log T +1.865 (18) = ---p T 6H To = 12770 + 3.0

*

T (cal/mol) (19) 6S To = 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.

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 NaZC03 [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.

-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 95vol-% MEK, 4.8vol-% SBA and 0.2vol-% 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/cm3 0.4519 cm3_/g 0.38 58.35 x I O-R cm 1.160 g/cm3

Analysis of their data shows that a mechanism of dual-site surface reaction is applicable over the entire temperature range studied (250-3100_C).

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.

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)

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/m3 ₎ (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

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 _w of a reactor tube is found by the expression [41]:

t

w

with: R - external tube radius p - pressure difference

over tube wall

(27)

(m)

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/m2 _• 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

_ .

__

. _ - - -- -- - -- - -- - - -- - -- -

-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

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

~

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-%

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