Design of 30.000 tonnes per year acrolein production

f

l~f~i

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Delft

Technische Universiteit Delft

" DESIGN OF 30,000 TONNES PER YEAR ACROLEIN PRODUCTION " MAY 1987

MR. F ARSHAD SALIMI

Faculteit der Scheikundige Technologie Vakgroep Chemische Technologie

1 -crnTmTS . / I. SlJMIvlARY 11. INTRODUCTION

111. DISCUSSION AND CONCLUSION IV. PROCESS DESCRIPTION

V. PROCESS KINETICS

VI. HEAT BALANCES

VII. EXPLOSlrn LIMITS

VIII. ECONCMICS

IX. CHEMICAL mGINEERING DESIGN

l . REACTOR DESIGN 2. GAS PREHEATER 3. REACTOR COOLER 4. PARTlAL CONDENSER 5. FLASH VESSEL 6. ABSORBER 7. STRIPPER COLUMN

A.

8. ACETALDEHYDE COLUMN 9. VACUUM DISTILLATION ~ 10.TWO-STAGE COMPRESSOR X. DISTILLATION COLUMN ( COMPUTER XI. START-UP PROVISIONS XII. HEAT AND MASS BALANCES. XIII. LIST OF SYMBOLS

XIV. REFERENCES XV. . APPENDIX •

- COHv1UTER PROGRAMS

- CERAMITC TOWER PACKING(NORTON CO. ) - DIFFUSIVITY CALCULATIQ~S ' )

...

3 4 6 10 12 18 I 22 Q2 33 33 · 33 34 36 38 38 38 42 42 43 44 59 60

62

• 2

-~ I • SLJrvMARY

This project is concerned with the design of an acroleinplant with a total capaci ty of 30,000 tonnes per year fram the catalytic oxidation Ol propylene. This process is a continuous one initially fed with propylene, oxygen, steam and inerts to a mul ti tubular catalytic fixed bed reactor. This plant is assLUned to operate 8,000 hours per year continuously. In order to proceed the reaction a bismuth-molybdenLUn catalyst is applied in the reactor. The reactor is operating at 6300_k and 2.5 bar. The overall propylene convers i on per pass is approximately 36 mole % and the acrolein and acrylic acid yields per pass are 35.8 mole % from propylene and 3.7 mole % from acrolein respectively. The main product, acrolein, is renoved from the reaction gases ( acrylic acid, acetaldehyde, water vapour, C02, CO, N2, and the unconverted propylene and oxygen ) first by the removal of acrylic acid-water solution in a flash column in which the inert gases ( C02, CO

&

N2 ) along wi th the unconverted reactants ( 02, C3H6 ), the main product ( acrolein ), and byproduct ( acetaldehyde) are removed as the top products. Further purifications take place in an absorber operating at atmospheric pressure to remove the unsoluble gases ( N2, 02, C02, COand C3H6 ) as the top and the main product acrolein completely

dissolved in water solvent is removed as the bottom product. The acrolein-water azeotrope is separated from the water absorption solvent in a stripper colLUnll operating at atmospheric pressure. The water solvent from the stripper collmll is cooled and reused as absorption solvent in the absorption colLUnll. Light product, acetaldehyde, is removed fram the acrolein-water azeotrope at thetop of the

acetaldehyde colunm operating at atmospheric pressure. The acrolein-water aze'jtrope free from acetaldehyde then undergoes further purification in a vacuum column oper-ating at 200 mm Hg in order to reduce its water content to less than I wt/wl %. The recycle stream taken from the absorption column contains unreacted C3H6 and 02 as weIl as the inert gases, mainly C02.. The excess products of exhaustive oxidation, C02 and CO, are purged off fram the system. In this plant a recovery system is not required ( also se~section1III).

11. lNTRODUCTION

Acrolein is a very toxic, inflammable liquid. The liquid vapourizes easilyand the vapours are readi1y flamnable in air at belween 2.8 anel

31.0 vol%. Acrolein vapours are twice as heavy as air. Ac)'olein polymcri7.:C:s easi1y and exothermally; therefore it is stabi1ized with 0.1 or 0.2

%

hydroquinone against radical-initiated p01ymerization, which Cá.n be cat:alyzed

by light ,air ,heat ,Ol~ peroxides. Acrolein should be stored and b~anspor't(~d

in the darl< under thc blanket of nitrogen at temperatuJ'es belo\\! 20°C, and i t should be used wi thin three months. The COlTTnercial production of acr"oJ" n by'heterogeneous catalyzed gas phase condensation Ol acetaldehyde and

formaldehyde was established by Degussa in 1942. Nowadays, acrolein is

produced on a large scale by heterogeneously c.::tlalyzed .C];as-phasc OX.lCl<lt:i.OI1 of propylene. Acrolein is an important intennediate [or numcl'ous sllbsté1nCC~'>'

The main use of conmercial,isolatèd acrolein is cUl~renUy l:he jwocluction

of

0, L-rnel:hionine, an essential amino acid used as él11 animul r eed supplement: .

In the production of acrylic acid, acrolein is not isolaled Il~Ol1l the gas -phase reactiun mixture but is oxidized further on a hctcrogencous catalyst. Acrolein is also used in polymer industry, especial1y for ,-lcl~yl(lni triL production , and production of glycerol. Acroleln has u bollil)(~ POillL of 52.7 oe and i ts water azeotl~ope' s boilil1,g point: is S2. 4°C at 1 at!l1. ACl'olein is markcled in the U.S. by the UniuIl Cé.1I'l>ich~ coqKlr':lti.ull \\'ilhil'l

thc [ollU\';in,g 1ill1i ts for the comnercial acroleil1. -Speci [ic gl~é1vi ty, 200

e

_120°C

-/\crolein,wl:.%, min.

-11 I of lCJ)~ solution in water ut 25°C ,max. -Total carbonyls other théUl acro1ein, wt% max. -Ilyclroqulnone, wt % -Water, wt %, max. 0.842-0.846 9H.O 6.0 1.5 0.1 - 0.25 1.0

\.

3

-With respect to the high toxicity of acrolein, the maximum allowable content of acrolein in the environment is 0.3 ppm. From the present discussion we approach three methods of acrolein productionj (1) via the pyrolysis of di-allyether, (2) the condensation of formaldehyde and acetaldehyde, (3) catalytic oxidation of propylene. The first two methods have been replaced by the catalytic oxidation of propylene, and this is due to the large ~1tities of propylene available fram the petrochemical industry, accompanied by great developments in research of good catalysts.

111. DISCUSSION AND CONCLUSION

For the choice of the process to acrolein format ion , one has to cOI.lsider important factors such as the kinetics of propylene oxidation which brings about the following discussion :

(i) Different types of byproductsj

(ii) Choice of acrolein recovery fram its mixtures with undesireable productsj (iii) The problem of disposing acrylic acidj

(iv) The choice of the recycle streamj (v) The choice of the used inertsj

(vi) Whether a C3H6 recovery system is necessary or not.

It appears that the choice of the catalyst and fram there the process choice for the acroJ.ein formation are not exactly the same as for acrylic acid. The flexibility of the types of catalysts give the acrylic acid manufacturers the selectivity in acrolein plus the selectivity in acrylic acid at the outlet of the reactor, while acrolein producers are not interested in acrylic acid since it is expensive to recover and a nuisance to dispose of. For this reason, we choose a suitable catalyst and the optimum operating conditions for acrolein production to bring the acrylic acid formation to a minimum in the reactor. In the kinetics section of this report, the detailed kinetics used for both acrolein and acrylic acid formation have been discussed. Fram this a suitable propylene ;

cOlwersion per pass wi th the recycle . for the pr'?~uction of acrolein

éf

the main product has been selected. Further, due to the development in the catalyst and increasing plant equipment etl'iciency, the acrolein production gets economically more attractive. As, nowadays, the high selectivity of the new catalysts ( bismuth~olybdenum metal oxides) towards the acrolein formation with the high propylene conversion per pass has made the acrolein production profitabIe. In this design, the pro pylene

and oxygen are the only fresh raw materials feedstocks . Steam lis also produced by the use of the water at 27°C fram the partial condenser, in the reactor cooler in order to save on the energy costs. Al ternati ve is the use of a steam generator in the acrolein plant to produce sufficient steam for the reactor as weIl as the rest of the system. Since the plant operates at relatively low pressure ; the , use of steam turbines could be an :advantage for the electricity production. The present design of this plant is based on the dilution by C02 produced in the reactor and steam at stationary state of the system. Furthermore, the system despite its purge, still could accumulate ~2 in the recycle stream and this could eventually lead to C02, N2, and steam as inerts. In this case, the maximum oxygen point in figure VII.l varies betweem 10 to 13.4 vol % since the max. safety percentage

of oxygen in C3H6 in the presence of C02 as inerts is 13.4 vol %. But the choice of the recycle stream, purge or even a recovery system depends on the costs. One idea was to apply air instead of oxygen as the fresh feedstuck to the system to

save on the costs of the fresh feedstock 02, but the propylene recovery system consisting of molecular sieves or propylene refrigeration seemed impractical and very expensive.

In terms of the kinetics of reaction, there have been several articles in propylene oxidation to acrolein. In this report, we have compared two types of kinetics towards

acrolein formation for the same type of catalyst. One gives high propylene conversion per pass ( 96 mole% ) with high acrolein yields of 88 mole %, while the other one

gives a lower propylene conversion per pass ( 35 to 40 mole % ) with the high selectivity of 98 % towards acrolein. The main differences in the two kinetics come from the diff-erences ir~ the catalys t composi tion, reaction temperature , the residence time and etc.

-- --- - - -- - - -

4

-The design of this plant is based on the proved kinetics of reaction, with low ... propylene conversion per pass ( 35 to 40 molè

%

)

and high selectivity towards

the acrolein.

. ~

For the 30,000 tonnes per year of acralein production, extracted from the econo-mic evaluation, it is concluded that this plant is attractive to invest on. On the other hand, one could consider several factors highly influencing the increase

in production capacity. First is the location of the plant. If the plant is sited to a cheap source of propylene, for example, near to a refinery or an ethylene plant and if the worldwide evaluation for the acralein market demand shows a positive sign then an increase in the production capacity of acrolein is recommended. For the location of the plant, one also has to consider the government legislation

over the tax and pollution contral. Further, the price of the raw material, propylene, highly depends on many factors including the oil prices and the market demand for

propylene and etc. IV. PROCESS DESCRIPTION

The fresh feedstocks , oxygen and propylene, af ter compression to 2.5 bar·.and 62.7 °C added with thp recycle gas ,containing the unconverted propylene, unconverted oxygen and inerts along wi th steam at 630 oK and 2.5 bar wi th the composition of

the gas mixture as C3H6 : 02 : inerts : H20

=

4 : 8 : 48 : 40 vol

%

are joined together in the gas mixer. The gas mixture, propylene, oxygen, inerts and steam then is

led to a gas preheater where its enthalpy increased via the heat exchange with

the reactor effluent. The temperature of the gas mixture then reaches 578 OK at 2.5 bar, suitable for the reactor.

The reactor has a tot al number of 7631 of 46 / 50 mm diameter tubes which are filled wi th the catalyst bismuth-molybdenum. The temperature along the tubular reaètor varies in the range of 578°K to 630oK, and Îurther the excess heat of reaction is removed from the tubular reactor by recirculation of the eutectic cooling medium, molten salt, which rise by 17°C from 578°K to 595°K in the shell side of the tubular reactor. The mol ten salt in turn ;Loses i ts excess enthalpy by the steam generat:ion in t:he reactor cooler ( H6). The hot reactor effluent: is immediately cooled at the outlet of the reactor by heat exohange with the inlet gas in the preheater ( H3). The gas preheater cools down the reactor effluent from 630 OK to 505°K. The further cooling of the react.ion gas mixture takes place in a partial condenser ( H8 ) by the heat exchange with the cooling water at 20°C.

The cooling in the partial condenser brings the reaction gas mixture temperature furbher down from 505°K to 333°K. The cooled gas mixture from the partial condenser is

led to a flash vapourizer operating at atmospheric pressure, from where the acrylic acid-water solution is separated from the remaining reaction gases. The gas mixture / from the top of the flash column is then pumped by a ~ump through a heat

exchanger ( Hll ) in order to cool do~n the gas mixture to 21°C. In this heat exchanger ( Hll ), the freon vapourizes in the shell-side at 5°C as cooling medium. The mixture at 21°C is then fed to a packed absorption column operating at atmosph-eric pressure lO order to dissolve acrolein-acetaldehyde solut:ion completely in water solvent or to eliminate the composition of acrolein in its mixture with the other g8ses ( also see mass balances). The off-gases from the absorption column, is partly recycled to the feed stream ( propylene and inerts ) and the remaining gases are purged off from the system ( also see section 111). The recycle stream contains propylene and C02 while the excess inerts along ( nitrogen and C02 ) with propane appear in the purge stream.

The aqueous solution of acralein from the bottom of absorber is then sent to a stripper col umn ( T14 ), d1ere it is stripped to give the crude acrolein; tr:.

bottom stream from this column iscooled via the heat exchange with the aqueous solution of acrolein from the absorber and reused as an absorbent .

5

-The crude acrolein from the stripper column is distilled in an acetaldehyde column to remove the low boiling byproducts, such as acetaldehyde as the top product. Then the crude acrolein is further purified from its water azeotrope in a vacuum column operating at 200 lTlTI Hg to bemove the excess water as bottom and 96 mole

%

distillate acrolein as the top product. The top product acrolein from this column contains less than 1 wt % water. In order to minimize the polymerization the ""hole system is stabilized by hydroquinone or a similar agent. Pipelines and apparatus are const-ructed preferably of stainless steel ( also see sections IX.7, 8 and 9 ).

Further, the fresh feedstocks used in this plant are taken as 99

%

pUr'ity which means that propylene and oxygen contain propane and ni trogen as impUr'i ty ( by 1 % ), respectively. In the stripper section an inhibitor solution of 0.02 kg / sec is also added to avoid the acrolein polymerization. This is a solution of 60 grams hydr'oquinone / Kg water.

'._"

-

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P I'o[)yl ene

+

;-.

- - - -R(!cyc.Le gast C:3i!6 ~ N2. ) Çl (19) .. CO]" Û"

()--

(~)

_ _ _ _ ... ~~ (i) ₍₄₎ ;)

1 bUl' & 21°C

OX)~en & nitrogen

... _ _ _ ---, • • (2) 1 bal' & 21°Ç Steam _ _ _ _ . . . ~ (3) sat.st:e 2.5 ba am at . r nnd 630 1.5 bar 1.0 bar Further st.8 at 357°Ç and st.9 at 232°Ç and st .10 at 600_{Ç an} st.ll

---_=

st.12

=========

d 1.0 bar ===:;::====== ========== Also st.14 at 21°Ç and 1.0 bal' st.15

========

=========== st.17

===================

replace H25 and H29 with V25 and V29 respectively. 1 bar & ~loÇ deg.K 2a5 baI'& C:2 62.7°Ç

0-

(6) 2.5 bar & 62.7°Ç n Ir

l'

:43

Z

r

_{. (9)} at 2.5 bal' (7a) and T mi>.:ture H4 (7b) (e) at 305°Ç

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: HlS CONOENSER H24

: VI6 PHASE SE PAR ATOR H25

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/ \ / ' / \ / REFLUX PREHEATER

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H28 REBOILE R H29 RECEIVER H30 PUMP V31 ACETALDEHYDE COL'JMN H32 CONDENSER REFLUX PREHEATER REBOILER VACUUM COLUMN " -CONDENSER

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REFLUX PREHEATER ACPOLEIN COOLER RECE IVER REBOILER Stream no.

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Temp. In'C

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-V. PROCESS KINETICS

The discovery that propylene could be oxidized rather selectively to acrolein marked the beginning of the current process of catalytic oxidation of alkene to aldehyde over the metal oxide catalysts. The BISl\1lJ1H-MOLYBDEl'JUM catalyst system was first discovered by SOHIO in 1957, which yield a fairly good selectivity of catalyst but still a low propylene conversion. Further, the kinetics of propylene oxidation on complex molybdenum-containing oxide catalysts have been the subject of many investigations. Depending on the chemical composi tion of the catalyst, i t interacts in different ways with the reaction mixture components, and different types of kinetics equations are observed. Fop. two-component catalysts containing molybdenum and bismuth, the kinetics of propylene conversion are described as first-order with respect to propylene and zero-order with respect to oxygenj the introd-uction of cobalt, cobalt and chromium, or cobalt and iron complicates the kinetic equations. Further for the design of this plant, the kinetics of reactions are extracted from the ref.(l). In this process, the propylene oxidation takes place on

a bismuth~olybdate oxide catalyst with added nickle, cobalt, iron, potassium, and

phosphorous, supported on silicon dioxide and containing 60

%

active mass. In terms of the phase composition, the catalyst is a mixture of molybdates of bismuth, cobalt, iron, and nickie. The sur~ace area of the specimen is 48 .1000 m2/ Kg, and the bulk density is 1 . 1000 Kg / m , ( M012BilFe2C03NilP2KO 2) ( also see ref.l ).

We can also list the complete propylene reactions sënemes as follows : C3H6 + 02 ... C3H40 + H20 (1)

C3H6 + l~ .02

='.

C3H402 + H20 (2) C3H6 + l~ 02 .. C2H40 + CO + H20 (3) C3H6 -+- 4~ 02 .. 3C02 + 3 H20 (4 ) C3H6 + 3 02 .. 3 CO + 3 H20 (5) Where for the reaction (1) we have

CH2 CH - CH3 + 02 ---_ CH2 = CH - CHO + H20 H

= -

340 KJ / mole of product.

r

The selectivi ty ~d th respect to reaction products and the rates of the partiai. reactions as functions of propylene conversion at 603 OK, are illustrated in fig.V.2 At the other temperatures, analogous curves were obtained. From the shapes of these curves, we can draw preliminary conclusions regarding the reaction schemes and their kinetic features. Acrolein and acetaldehyde are the products that are formed in parallel from propylene. The initial selectivity with respect to acrolein is quite high ( 90

%).

The initial selectivities and the rates of accumulation of acrylic acid, CO, and C02 are very low, indicating consecutive reactions in forming these products. The acrylic acid is obtained from acrolein. The drop in selectivity for

acetaldehyde with increasing conversion indicates the post-oxidation of the acetal-dehyde. The CO and C02 may also be formed from acrylic acid and acrolein. Al though, the excess amount of CO, C02 have no effect on the reaction.

W 1 C3H6 - - - -. ~C3H40 W 3 . . C3H402

I

W 5

~I

W 4

-~

W 6 W 7 ~ C2H40 ~ CO,C02

.-

;:-...

_~ (J Il)

-

Q) ti)

W

.10

9,

mOles/m

2 •

sec

20

10

1,0

fl,?

lOt}

ófJ

7

7

-W

.10

9,

moles

/1lJ2.

sec

25

/.5

611 ~

20

~

c

o ~

'~t

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1

Z[J

40

50

80

!O{j

A,

nlo

Fig.V.l Fig.V.2

Fig. V.l Rate of overall propylene convers(on (1), rates of accumulation of acrolein ( :2) ,

acrylic acid (3), products of exhaustive mudation (4), and acetic acid (5), '-md

:- ·sëlect:ïvities with respect to acrolein (6) • ac rylic acid (7). acetic acid (8), and products of exhaustive mudation (9) as functions of propylene conversion level.

Fig. V. 2 Overall propylene conver'Sion rate (1) and select-ivi tj es wi th respec.l: to

acrolein (2), acryl~c acid (3), acetic ·acid (4), and sum of products of exhélustive

mudation ( CO + C02 ) (5) as functions of propylene con'v"cr'sion, wi. til followj lig

" I _. \

I

~

\

I

I

t

I

I i :

s

-Further, as the concentration of the water vapour in the reaction mixture _{the acrylic} is increased,

acid _{and aceta1dehyde rates increase, the CO and Co~ aCCLUnul<1tion}

rates

decrease, and the acrolein accumulation rate increases slightly. The products of

exhaustive oxidation are obtained by three paths. In the oxidation of acrolein,

an increase in the water vapour concentration accelerates the reactior \s of CO + C02

formation . The tot al rate of acrolein conversion remains unchanged. The presence of water vapour lowers the rate of post-oxidation of acetaldehyde. Thus the depen-dences of the rates of the partial reactions in propylene oxidation on an

eight-component catalyst are described by equations that are first-order _{with respect to} the product being oxidized as follows.

Kl . C .( 1 + K_{2 CA} ) W . - W

=

_-.:~_..!p:..:r=--_{_ _ _ _ .:::...!.~c:::.!ro~=-.}

Acroleln - 1 (1 )

1 ₊ K C

3 Acrol.

\~en _{the reaction is carried out in an excess of water}

vapour _{which is} of ten the case, and in this case, the kinetics equations are as follows:

\V

=

_{K C}

Acètal. 4 pr + K C 5 Acro - K 7 C Acetal. (2) = W

2 +\'>'5 - \'1 7 W _{Ac rA} = K C _{6 Acrol.} - K C 9 AcrA ₍₃₎

₌

W 3 - W 6 W_{CO ,C02= K7 C} + K C + K C (4)

=

W 7 + W4 + W6 Acet . S Acrol . 9 AcrA

\~ere. \'1 _{is. the. rate of formati?n, C is the concentration and finally K. 's}

reactlon lunetlcs constarlts as listed below along wi th the activation 1

energies,Ei " Values of Kinetics constants are at 603 OK. TABLE V.l

Constants (exc~pt _for

i Lit/m2 . sec. ;, K

2 . & K3

E, KJ'/mole _{Constants E,KJ/mole}

tin )

in Lit/m2_.sec

Kl 1.20 .E-6 29.3 _{K6 0.66 E-6} 40 ')

K

2 22.5 E+3 Lit./mole 46.1 _{K7 5.00} E-6 91.3

K3 1.20 E+3 Lit./mole _20.9.

KS 0.20

.

E-6 71.2

K4 0.17 E-6 _44.0

K9 "0.60 E-6 e3.8

KS 0.10 E-6 41.9

Where W Acro' W AcrA' Acet' W and Wco,co:--~ .. · ar. e the respective 1"ates oÏ formati.oll of acrolein, acrylic acid, acetaldehyde, and products of exhaustive oxidation; C ,CA '

C C C C _'

-- -prof) crQ

Acry' Acetal. ' .H2Q..' CO.C02 are the respective concentrations of propylenè, àcr'oleln,

acrylic acid, _{acetaldehyde, water vapour, and finally tr.e products}

of exhélust i ve . oxidation; K. _{'s are the reaction kinetics constants at 603 OK.}

9

-The reaction mechanism in the formation of acrolein, which is the main product from propylene oxi dat i on , has been studied by many investigators. The limi ting stage

OF reaction is the dissociative adsorption of propylene, accompanied by the abstraction of a proton and the formation of a '" - allyl c~plex. Abstraction of the proton

takes place on the nucleophilic oxygen of the catalyst. Subsequent interaction of the ?t -allyl complex with the electrophilic oxygen leads to the formation of acrolein. In our case, the acceleration of the acrolein formation reaction as the acrolein concentration is raised may be due to an increase in the number and strength of the nucleophilic centers.

Acrolein is an electron-donor reactant. The electron density in its molecule is displaced toward the carbonyl oxygen. On molybdenum-containing catalysts, species of weakly adsorbed acrolein have been identified, formed by the interaction of the

carbonyl oxygen with a high-valence metal cation. We believe that when such interaction takes place, there is a redistribution of electron density on the catalyst surface, strengthening the nucleophilicity of the oxygen bound to these centers, and this should facilitate attack of the C-H bond and abstraction of a proton. Thus, in the course of reaction, under the influence of the acrolein-containing reaction mixture, a modification of the catalyst surface takes place, leading to a change in the state of the surface oxygen and in the rate of the acrolein formation reaction.

I

..

10

-For the calculation of the heat balilllces over the acrolein p1ant,Cp values

(specific heats) of gases have been obtained from the fo11owing equation:

Cp = a -;- bT +cT2 + dT3 (1) in J / mo1e OK

Then aîterwards a correction was app1ied for the non-ideality of gases at a

cO!~responding reduced temperature and reduced pressure . The corre1ation used for this correction was pub1ished by the Edmister. The speci:fic heats of

other gases( steam,CO,C02 ,N2 ,02 ,air) along·Jwith other thermodynamies prope-rties have been obtained from the ref.(ll) and other references.

TI1e specific heat mean values and heats of formation of gases are tabulated

in tab1e ( 1 ). The specific heats of 1iquid aCl~olein and water are also 1isted be1o\'.'. The constants for the equation(l) ill1d other properties of the compounds are l i s ted in bab1e (2) .

. Compou-Hf Cp values of gases in KJ / Kg OK H v nd KJ/mole 294 323 . 343 370 423 57ts 603 650 673 KJ / Kg at B.P. C3H6 7.6 1.51 1.53 1.7 1.81 2.0 2.494 2.51

-

2.8 437.774 C;~r-i8

-

1.7 1.8

-

2.0 2.16 2.85

-

.

-

3.18

-02

-

0.92 0.92 0.93 0.94 0.95 0.995 1.003 1.02 1.030 213.25 N2

-

1.04 1.04 1.04 1.042 1.05 1.07 1.075 1.086 1.086 199.23 H20 . _{246 1.864 1.870}

_-

_{1.89 1.91 2.0 2.015 2.05 2.10 2258.3} C3H4G - 74.6 1.22 1.30 1.32 1.40 1.53 1.973 1.927

-

2.044 505.6 , C2H40 - 167 1.25 1.32 1.35 1.43 1.60 1.9045 1.935

-

2.10 587.5 C02 -- 394 0.134 0.871 0.890 0.920 0.95 1.061 1.075

-

1.105 390.14 CO - 13tl 1.04 1.04 1.040 0.910 1.051 1.081 1.087

-

1.110 215.85 C3H402 1.068 1.118 1.680 1.721 639.1 Air 1.005 1.006::

-

1.01 1.017 1.0454 1.051 ·1.063 1.070 213.54 Tab1e VI.1 _{.. . ..} .

Canpound

T

P

B.P.

t-r(.~ •

a

b c

d

c c

OK

_Bar

_°c

_Cr.

_-

_-

_-

-.

C

3

H

6

365

46.0

-47.8

42

b.710

2.345E-1

-1. 16E-4

2.204E-8

C3HS

369.8

42.5

-42.1

44.1

-4.224

3.062E~1

-l.586E-4

3.214E-8

C

3

H

4

O

506.0

51.7

52.8

56.06 11.97

2.15E-1

-1.070E-4 1.905E-8

C

2

H

4

O

461.0

.

55; 7

20.4

44.054 7.716 1.82E-l

...

-1.006E-4

2.380E-8

H

2

0

"

.

647

.

.

3 220.5 100.0 18.0

32.243

1.

923E-3

1.055E-5

-

3

.596E-6

CO

.-

132.9 35.0

-191.5

28

30.870 -1.285E-2

2

.79E-55 -1.

271E-8

CO

2

204

.2

73.8

-7

8.

5

44

,

19.8

7.343E-2

-S.601E-5

1.715E

-8

1'2

126. :2

33.9

-195.8

28

.0l

31

.15 -1 .

356E

-2

2

.

679E

-5

·--1 .1 68E-8

~

_{- ._._} . . _----

-

154.6 50.5

-

_- -... ._.

-183.0

32.0

·

28.166

-3.68E-

6

1

.745E

-5

-1 .0

6

5E-

8

Air 132.3 37.2

-

28.9

-

-

. .

-

-C3H402 615.0 56.7 140.8 72.0G '1.742 3.19 E-1 -2.352E-4 6.9-:-5E-8 T."'.BLE (2)

r

I~

I

I

I

I

I

I

I

I

i

I

I

I

I

I

I

I

I

1 11 1

-The specific heat va1ues of 1iquid acro1ein and water are as fo110ws H20 : 4.2 KJ 1 Kg C3H40 2.14 KJ 1 KG

The molten salt bath used in the reactor cooler has the fo11owing properties:

,,=

0.36 W 1 mO C,

f

=

1838 Kg 1 m3 Cp = 1.56 KJ/KgOC, '"

'v=

2~.10 -3 NS/m 2

The feed gases are introduced at 1 bar to a two stage turbo compressor in order to bring the pressure of the gases to 2.5 bar. The compression is a non-isothennal one and the gas temperature af ter compression is calculated by the following thennodynamic expression

' ( - 1 / ' (

T2

1

Tl

= (

_P2

1

PI)

=

Hence T?

=

~94

. ( 2.5 1 1.0 )0.145 T

2

=

335.77 OK or 62.77 °C

The outlet gas fram the preheater is at 305°C.( preheater, H3 ).

The heat of reaction is calculated from the values of heats of fonnation as 18,092.0 KW.

For the heat balanee calculations, the change in enthalpies due to the change. in acrolein phase anct water phase in th~ s~tem, HL ___ G have been taken lnto account. ïhe latent heat.. of vapourlzatl.On of gases are found from the literature at their boiling points and then they are converted to the desired temperature by the following equation fram WATSON.

Where n

=

0.38 1 1 Tr2 )n T rn T _{r_} ? = Reduced temperature at tne desired temperature.

T

=

Reduced temperature at the nonnal boiling point.

rn

The principle calculations of the heat balanees were based on the following relationship T HT =

A

Hf +

J

Cp dT +

h

H 25 Enthalpy at 1 atm, m1d temperature , T cC.

Cp

=

Specific heat. é::.. H

=

latent heat fonnation at 25°C m1d 1 atm.

Standard heat of

fonmation at. 25°C

and 1 atm. Hü\\'ever, enthalpy of the streams \\nere there is not m1y !'ea·.::tion involved, is calculated by the elemination of the enthalpy of fonnation of the mixture compounds. SA the total enthalpy of each stream is described as follows :

And the heat of reactiün for the reactor design is calculated fram the following

equation

~6Hf

- reaetants

= ~H

..

12

-VII.EXPLOSION LIMITS :

Propylene is very inflarrmable and explosive in the presence of oxygen. In

the design of this plant, we are to investigate the explosion limits at four area in which the inflammable compound, propylene, is in contact with oxygen.

(i) Tubular reactor:

Although, the tubular reactor is operating non-adiabatically and the heat is being removed continuously by the eutectic cooling medium, molten salt, there are still sufficienb. safety precautions required to ensure the safety of the plant in terms of the explosion. From the mass bal anc es , we can see that the feed to the reactor contains propylene, oxygen, steam, N2, and a large quantity of inerts ( C02). Steam added to the feedstock is to keep control of the reactor temperature , and also act as a diluent. At 20°C and atm. pressure extracted from the Gasunie tables ( see ref.8 ),

LEL VEL

C3H6 in 02 2.1 vol % 53.0 vol %

air 2.4 vol % 10.3 vol %

Table VII.l.

Maximum safety percentage of 02 in propylene C02 as diluent

Propylene 13.9

N2 as diluent

11.4

Feedstock to the reactor has the following compositions

C3H6 _{2.285 kg sec} -1 _{54.405 moles sec} -1

02 _{3.483 108.84 ,} C02 28.7245 652.83 H20 9.79155 544.0 N2 0.03485 1.2446 C3H8 0.02285 _{0.52/ total 1361.8436} -1 moles sec 4 % 8 % 48 % 40%

In order to calculate the vapour densities and from there the volumetrie flowrate

of feedstock to the reactor, we apply the following equations

v

=

Z • R • T / P ( 1 )

And then using the Lee-kesler correlation for the calculation of Z ( the

compress-ibility factor) and using the equa~~on ( 1 ) expressed dimensionally as

M • P

10.72 . Z . T Where

( 2 )

r

-3

= density

in lb/ cu.ft multip1y by 16.02 to kg m

v

Mw = Molecular weight

P pressure in psi

We are operating the reactor non-adiabatically at 2.5 bar; P = 2.5 bar or 36.23 Psi

..

-3 = 4.10027 kg m

I'

=

2.982 kg m -3 v 02 3 V C3H6

=

0.7874 m 13

-f

_v

₌

3.914 kg m -3

f

₌

C3H6 vH20

P

₌

4.10027 Kg m -3

I'

₌

v C3H8 vN2 3 V 02

=

1.5748 m 3 V C02

= 9.44336 m

-3 1.6774 kg m

.

2.61 Kg m -3

.

3

=

7.871 m V 3 N2

=

0.018 m VC3H8

=

0.007512 3 m; V 3 total

=

19.70212 m /sec. Further, the gas mixture could reach to 6300_{K or 357°C, so again we have :}

P

₌

3.53136 -3

f

₌

3.371 -3

f

₌

-3 v C02 Kg m v C3H6 Kg m _v H20 1.444 Kg m 2.5682 -3

₌

-3

f

₌

-3 fv

=

Kg m

..

f

3.5314 'Kg m 2.247 Kg m 02 vC3H8 vN2

We now consider the explosion limits equations ( also see ref.8, P.IIl.24 )

Eb=

E,

E,t 0 100 E m

=

100 100

---ZX.

/

E. . ~ D,~ ( ₁

₊

_---

t

-

20 ) 1300 ( 1

...

_---

t

-

20 ) 1300

Eo

=

LOVier explosion limit

.

_Et

~

N. =

%

component i, ( excl. 02 1 ( 3 ) ( 4 )

( 5 ) ( Applicable for

UEL

at 't' temp. ( 6 ) ( Applicable für La at 't'

temp. ( 7 )

~ = Upper explosion limit

N.

=

100 )

1

E. Lo\\er / upper e;-.,:plosion limits of component i

1

E Lo\\"er / up:'Jer explosion limits of the mixture.

m

F'rom ther~ .. nd with refcrence to table (VII.2) , we have

t

( ₁ 305

-

~O ) _UEL

Et, . =

+

--- 53 = 64.6~ at 305°C

,1

1300

From eqt.lations 6 & 5

t

Eb,i

305 - ~O ) . ~.1 1.64 LEL at 305°C

14

-Now having these values and with reference to table (VII.l) and ref.(8) we have the following LEL and UEL at 305°C in " OXYGEN" for the systern of propylene-C02-steam ( inlet to the reactor ).

E.

= 100 /

=

70.73 vol %

Equations 3 and 4 .

Et,

=

100 / 91.36 = 1.79 vol

%

1.64

Using the same method as above for the " AIR " we then have

Et

b = 12.56 vol

%

Eo = 13.75 vol

%

~

= 1.874 vol

%

_Eb

= 2.05 vol

%

Now for the volumes at 357 oe or.630 OK, we have the following :

V e3H6 0.678 3 V 02 1.3486 3 V H20 = 6.7781 3

=

m

₌

m m V e02 3 3 3 = 8.13412 m V N2 0.01551 m Ve3H8 = 0.00647 m V total = 16.9607 · t EL. .

= (

1 + ~,l 357 - 20 ) . 53 = 66.74 vol

%

UEL at 357 oe t E.~_-a,l . =(l+ 1300 357 - 20 ) .2.1 = 1.5556 vol

%

LEL at 357 oe 1300

Now we apply the same method as before to obtain the following

Ei:,

= 100 / 91. 91 = 72.26 vol %

66.74

Et,

= 100 / 91.91

1.5556

= 1.692 vol

%

From the infonnation obtained the triangle is: ',constructed for the detennination

of the explosion area. As the first approximation, it is of ten stated that the

lower explosion limit is detennined by the concentration of the inflammable gas

independent of 02 and the LEL is to a certain extent constant, while the UEL

is detennined by the 02 concentration independent of theinflammable gas concentration.

Below the lower explosion limit, there exists too poor a gas mixture while above

the upper limit ( UEL ) there exists too rich gas mixture. It is also importa~t

to consider the LEL as a criterium that plays a role in the dilution of the gas wicn

the diluting agents or inerts, in order to obtain a poor gas mixture. For the

calculations, the effect of H20 as diluent is taken the same as e02 or N2. The

effect of pressure on the explosion limits (at 2.5 bar ) are not considered

in explosion limits calculations due to operation at relatively low pressure.

Further, for the s?;raphical interpretation of the explosion limi ts, we have still

used the ' air line '. This is simply due to the fact that the amount or the' accumulation of N2 in the systern could vary. Secondly, in all cases, the gas

composi tion point is not effected by the posi tion of I-he ' air line '. Consequently,since

the ' air line ' and the ' air composition ' only simplifies the detennination of

15

-(ii) Af ter the tubular reactor

Si'nce the concentrations of the inflarrrnabie gas, propylene, and 02 in the reaction mixture are reduced then the position of the inflarrrnabie gas in terms of explosion improves. In other words, the reaction mixture stays· outside the explosion area

af ter the tubular reactor. .

(iii) Recycle stream :

The recycle stream contains the unconverted propylene, oxygen and inerts ( mainly C02 ) . For this purpose, from the mass balances we have the following

f

= 32.58 Kg m -3

_f

= 34.13 Kg m -3

f

= 24.83 Kg m -3 v C3H6 vC3H8 v02

f

= 21.723 Kg m -3

f

34.13 Kg m -3

f

= 13.961 Kg m -3 v N2 vC02 vH20 3 -1 V 02 = 0.09025 3 -1 V C02 0.8416 3 -1 Vtotal =O.~6~

V_C3H6 0.0357 !il sec m sec = m sec

recycle m s ~x. = 90.68 vol % 1 C3H6 t

Eb

,

i

=

3.7 vol

% ;

( 1 + 21 -20 1300 02

=

9.32 vol

%

53

=

53.0 ) . 2.1

=

2.1 = 100 / 90.68

-53--

= 58.4 vol

%

C02

=

86.98 vol

%

= 100 /

=

2.316 vol

%

(iv) Recycle stream plus the fresh feedstock, propylene:

Al though, the composi tion of p~ 'Jpylene in l:he gas mixture increases, but the posi tion of the gas mixture remains outside the explosion area. This is compared with the position of the recycle stream in figure VII.3.

C3H6 = 7.27 vol

%

~x. 91.17 vol

%

1

02

=

9.0 vol

%

inerts, C02

=

83.9 vol

%

The explosion limits here also remain almost the s~e as the case (iii), the recycle stream.

- 16 -Inflanmable comp·. Incrts in vol %

·

L

X. Ratio E."-plos i 011 _{1 imits}

Component Vol % C02 & N2 Steam 1 inert linfL.lRl _~,i

able I'b,i

C3116 4 48 40 92 22 1. 79 ·;().73

02 8 8 2.05 I;l. 75 _{IN AIR}

(a) 5 ys tem at 578°K. ( Inlet to the tubu ar reac tor

Inflanmab1e comp. Inerts in vol %

~Xi Ratio Exp1osion llmits

Component Vol% C02 Steam inerts _rrmab1e / inf1a- _{E I~}

O,i b,i

C3H6 4 48 40 92 22 1.692 7:2.26

02 8 8 -

-(b) System at 630 oK, ( in the tubular reactor ).

Inflanmab1e comp Inerts in Vol %

LXi Ratio Explosion 1 imits

Component Vol% C02 Steam _rrmable inerts/inf1a- E _{O,i ~,i}

C3H6 3.7 86.98 - 90.68 23.5 2.316 511.4

02 9.32

-

- 9.32 2.3 10.5

(c) System at 294°K, ( in the recycle stream ).

Inflanmable comp. Inerts in Vol %

'2'

X. Ratio Explosion limit!:

1 inerts linfla

Eb,i

~i

Component· Vol% C02 Steam rrmable

C3H6 7.27 83.9 - 91.17 11.54 2.32 !ïH

.-1-02 9.0 - - 9.0 2.3 Itl.:;

(d) ."ystcm at 294°1{, ( Lhc I'ccyclc strcam [11us I"'esn I'roI'y.i.e;nc fcedstock )

~

Figul1

e VU. l: FLA\1\1Al31LITY DIAGR4'VI FOH

THE SYSTEM

PROPYLENE-OXYGEi'l-NITROGEN-STEi\M AT 2.5 BAH

Al\lD 305°C.

13 FEED i\IIXTUHE PUlNT.

VOL% C3116 INEHTS VOL%

\

o

~

9

0

I

~vr

\ (

_o

\ ,

_v

_{\ /}

_\vl

_\v

/

_.

_\

I

I

~

I

~~

V

v

-

___ _ .. 100 10 0 -'~.' . 100 90

eo

. ... 70 . .-.. . .. " ... . 60 50 40 30 OXYGEN VOipo --' '-I PJ

Figure VII.2. FlamnabiliLy diagram for

the syst-em

propyltlle-0xYgen-C0'2- ni rrogen-! steam at

i

2

.5

bar a.nd 357 °cI. Propylene Vol%

..

I C3H6 I _I

i

i

I.

I ! I

I

,

I

r

'

erts,1 Vol

t

%

, I /

40

'

.

.

\

: i

I

1

I

, 1 1

I

~~~

(

I

I

.. / i

i

i

.

.

/

I

I

I i

I

i ! .. 1 . . \ . I ,

i

i

I

i

I

.. I I I I I I

I

I I \ I . " 1

~

I I .. I I I i

I

.

I

I

I

i

!

I

I I

I

·

I

I

I

, I I . . 1....

I

..

!

I

--

1

-

·

-

I

- :

i

'

. i

Fiy,ure VII.3. Flammability diagram :for the

systerr. propylene

-oxygen-C02-nitorgen at 1.0 bar and 21°C.

1:3

1 8

2

Position ol' Lhe r'l:!cycle stream. Posi tioll Cd th.· recycle st .'eéU11 i

plus the fr'esh reed propylene. :

•

C3H6 ' 20

I

I

I

i

i

i I I , ~[NERTS VOL

%

I

I.

I : /

S;t;.

:

'

/

~ ~ '

'(-

'~/ . ,~ F o

, z ' v

L

~Ssf

\ t

\/~

\

_I i ' \IJT

\\7

~

"'\

~--H

i

~

;

LOb

W() 90 HO 70 60 50 40 30 20 10 o OXYGEN VOL

%

I

I I I ! I, j

I

" I

I

I

I

j,

!

_i

I..

~ n

18

-In order to evaluate the investment cost of the acrolein plant, the model published by TAYLOR in 1977 is applied as follows. First of all the table for the costliness index is constructed based on the significant process steps of the acrolein plant.

C

I

B

=

45 f . pO.39 . ___ ₁₇₅ 1_ in which

N Si

f =~ (3,1.) the costliness index

=

_43.5

1

Cl = the cost index in the U.S. for the yr. 197B

eq.(l) Dec )

₌

p = The capacity in Kt / yr, for the acrolein plant '= 30

191 ,Kt / yr. I B = 45 . -43.5 30 0.39 191 = ' Kf. 8,049.601' Ol' . 175 f. or $ 8,049,601.45 12,074,402.19 (TABLE VIII. 1 )

THROUGH MATERIAL REACTICN PRESSURE TEMP. OillER TOTAL PUT OF CONS- & SCORE

TRUCTION STORAGE TIME STORAGE / , .-HANDELING

---C3H6 1 1 0 0 1 1 (expl. ) 4 AIR 3 1 0 0 1 1 6 RECYCLE - 5 1 0 0 0 1 - 3 ACROLEIN 0 1 0 1 0 l(toxic) 3

---PROCESS

---l)CCMPRESSOR 4 1 0 0 0 1 6 2)REACTOR 4 l -3 0 1 1 4 3)QUENCHER 4 1 0 0 0 0 5 4)ACRYLIC ACID 'SCRUBBER 3 1 0 0 () _{0 4} 5)ACROLEIN ABS-ORBER 6 1 0 1 0 0 8 6)STRIPPER 6 1 0 0 0 1 8 7 ) ACETAlDEHYDE COLUMN 0 1 0 0 0 0 1 8) VACUUM COL-UMN 0 1 0 0 0 1 2 \

For the calculation of the operating casts, a simp1e model extracted fram the Chemische fabriek (ref.8) is used. The model is as :~l~GWS :

-( K

=

K + K' + K' T P I. L (2) COSTLINESS INDEX 2.8 4.8 0.4 2.2 4.8 2.8 3.7 2.8 8.1 8.1 1.3 1.7

-43-:-5

We consider the defini t i on of the product (p) and the

we have;

of the K , consisted of two elements the quantity cost / P quantity, in general per tonnes (K ) so

p

K

=

K P (3)

p PO

And the K is the Slm ol ~he fresh feedstocks per ton. of the product, thus;

p

K

=

~ v . . q. Where v. is the cost /ton. of the fresh feedstock. (4) P i ] . ] . ].

I

-- 19

-q, is t;he quantity per ton. of the product necessary.

1 I< a. I< P ( 4) p p K' I f I (5 ) K' = d L (6 ) L

por thc calculation of the "L" the labour costs lor the acrolein plant \\'e apply the 11 STEIN REU\TIONSHIP".

E : 7.0 (7 )

E no. Ol employees = 63 for acrolein plant.

1 = Inves tment in i'vl$

Labour cost per hour in U.S. is $,8.5 in 1978

Total labour custs fOI' the 30 I<t /yr is $ 4,284,000.0 for 8000 hrs / 63 men.

Now'for thc calculation of I< , we take the price of C3H6 = $ 428.57 / ton'. Take the exchange rate for p L / $ = 1.5 andf:' / j ': 3.5 and then $/j = 2.33 .

The only raw materiaJs that we purchase for this plant are :propylene and oxygen.

Now we consider the equations of 2,4,5 and 6 ; with the help of the table

11.12 of rel.8, the values for a,d,f according to "Best model" are 1.13 ,

2.6 and 0.13 respectiv21y.

KT = a. Kp . P + f . I + d . L (8)

KT = 1.13 . ( $ 17,930,138.33) + 2.6 ($ 4,284,000.0)+ 0.13 . ( $ 12,074,402.19)

=,'$ 32,969,,128.,6 TOTAL PRODUCTION COSTS

VARIABLE OPERATING COSTS

Alternative method for calculation of I< is from the fol10wing equatiun:

,p

K

=

1<0 ( 0.01 . E + 0.003 I< + 3. S - 2 S + F . st

P v P

In which : E

=

K\Vh / tonne of prodLlc t

I<

S :

Cooling water / tonne of pI'oduct Tonne s team used / tonne of product

j / tone of prod~çt.

(9)

SV =

P 10 :

Net steam produced ( tonnes ) / tonne of product st=

Sensible heat carrier m3(torule) /3torule of product Calorific value in 10 and is GJ/m (tonne)

constant ( f 15 / GJ )

Note that the price of oxygen feedstock is taken as f 0,30 / m3 and the total . cost of this raw material, per year is $ 3,945,727.8. The exchange rates mentioned

above are also subj ect to fluctuations and these valües haVE: been rOLlghly as ' above for' the first six months of the year 1987.

Cost of raw propylene / year $ 13,984,410.53

Cost of feedstock oxygen / year $ 3,945,727.8

20

-FOI' the calculation of

I'

E " the work of compression W"lder the isentropic

condi ti on for compressors (Cl, C2) is calculated as 172 KJ /[{g(which is 1.711 1011 . KJ /yr.

and then E = 1584.2 for 3,O,000tonnes per year of acrolein product.

F

₌

1.0715

S

₌

2.45

P

S

₌

11.1046

v

K

₌

312.21 Only for cooling water excluding the used refrigerant 12.

E

₌

1584.2

st

₌

37.68

K calculated

p from eqn. (9) with the use of above mcntioned data is $ 550.87 /tonnes

of product acrolein.

For the estimation of HOI, we assumcd the capital is borrowed at 15

%

interest for the duration or the project, and for the tax calculations,

we tal<e straight line depreciation of 10

%

per year for the acrolein plant.

We asswne the plant is operating at its full capacity and the total acrolein

produced ean be sold, then we have the fol1owing calculations :

Inves tmen t, fixed

, working capital

Total funds commi tted

Armual receipts from sales Cost of sales

Receipts minus cost of sales

Depreciation, 10

% /

yr.

Net taxable income

Income taxes, at 50

%

Net income

Now \\'e apply the following equa:t~on for ROl :

$ $ $ $ $ $ $ $ $ $ 12, Q74,402.19 5,000,000.0 17,074,402.19 28,500,000 17,930,138.33 10,569,él61.67 1,207,440 .219 9,362,421.451 4,6811 310.726 4,681,210.726 ROl . 100 % ROL $ 4,6RJ ,210.726/ $ 17,074,402.19 o

27

.

4:10

21 -ECONOMICS

For the acrolein plant of 30,000 tonnes per year capacity, a crude estimation

of the profitability of the plant has b~en examined. The total capital cost of

the plant is estimated through the use of the " TAYLOR MODEL ", I

B = $ 12,074,402.19

based on the cost index of 191 in Dec. 1978 in U.S.

The total operating costs of the 30,000 tOIlnes per year of acrolein plant

accor-ding to the simple model in ref.8 is estimated as $ 32.969,128.6 (variable +

fixed operating costs). The variabie operating costs calculated separately as

$ 550.87 / turu,es of the acrolein product ( see ref.8). In the calculation of the

operating casts, the costs of the refrigerant 12 is not taken into account. In

gene-ral, the variabie operating costs should be by a factor of 15

%

higher due to

neglecting of the pumps operating costs and other essential utility materials.

However, with the cost of capital around 15 % and inflation (1979 rates)

at about 10

%,

giving a minimum acceptabie rate of return of about 25

%,

the ROl

af ter tax of 27.4

%

(as a crude estimation ) seems attractive. Further, the

location of the plant and the economie positiol, of the plant is of vital importanee.

At the present, the inflation in Holland is around 2

%

,and with the interest rate

of 8

%,

giving a minimum acceptabie rate of return of about 10

%,

the plant is

quite attractive. Furthermore, the economie stability of the land ~here the

plant is to be instalied and the crude oil prices can significantly effect the profitability of the plant during its life time operation.

22

-IX. CHEMICAL ENGINEERING DESIGN

1 • REACI'OR DESIGN

(a) 'lbeory behind the plug-flow reactor

'lbe concentration and convers ion per pass of the propylene along the tubular reactor changes continuousl y • For this purpose, we introduce the following concept for the tubular reactor; for propylene across the element, we can draw the following rnass balances :

Q.C 1= Q . ( C 2 _{p, p,} +

d C 2

_{__ EL_}

dL

L + ( 1 )

Now, since the concentrations of the acetaldehyde and the products of exhaustive oxidation are rather low, consequently, we could apply the following equation;

( 2 )

Further, in order to develop the convers ion in an ideal tubular reactor where the reactant " A " is converted according to reaction

A

---I.,.

Products

From a molar balance of " A " over an infinitely thin slice of the reactor ( fig .IX.lb ),

we can derive the following

Out

=

In Conversion - Hold-up

( - rA) . F dx - 0 So _dCA C/J v

----

.

dx

=

-

( - r )

.

Fdx dx A dCA ( ) F --Or

(i-x--

=

-

-

_rA

.

_~~- ( 3 )

With the initial condition

at x

=

0 ( 4 )

If the rate equation is

known

and relatively simple, the concentration profile along the reactor length and so the concentration at the end of the reactor le-'1qth can be calculated analytically from eqns ( 3 ) and ( 4 ). In the case of the

design of the reactor for the acrolein plant, we IlEke use of the finite differenc2 ~ .. ~=,::._~

method. of Runge-Kutta and a basic program set up in ref. ( 24 ). For this reason, the djroensionless quantities have to be introduced; such as the dimensionless reac+-'..Jr length Z

=

x / L and the degree of conversion ~

=

(CA - CA) / CA ' then

equations (3) and (4) become : 0 0

_9~

( - r A ) ,

re

= f ( ~ ) (5)

=

-CAo

.

dZ s

With the initial condition

- 23

-(b) Autothennic fixed bed reactor concept:

The principle reason for the autothennic behaviour is the effect of the feedback; the rna.ss entering the reactor attains ilnmediately the reoction temperature • It feels the effect of the heat of reaction. A situation wJ th multiple steady states can also be obtained in the catalytic fixed bed reactor. There are several ways to realize the feedback. It can be done by preheating tlle feed in an external

heat exchanger with the reactor outlet stream or al ternat i vely to design the tubular

reactor as a heat Jexchanger and if appropriate one can apply bothmethods.

(c) Multi-tubular fixed bed reactor ( Chernical engineering design) :

The catalytic propylene oxidation reaction is highly exothennic and the excess heat of reaction is being remaved continuously fram the non-adiabatic reactor by the cooling medium, rnolten salt. The heat of reactioll is calculated fram the heat of fonna.tion of gases listed in table (VI.l) . The calculation of thenecessary heat to be removed fram the non-adiabatic reactor is ~s Follows

Q - Hr

=

H8 - H7b

P

= ,:

,..

Ilr~)

-

,.

( 7 ) Where H ct' = 18, 091.5456 KW rea J.on H8

=

76, 641.876 KW at 630 0 K H7b

=

67, 364. 0835 KW at 578 0 K Then Q

=

18, 091.5456 9, 277.8 to be removed

=

8, 813.75 KW

The cooling medium used to remove the excess heat of reaction fram the tubular

reactor is rnolten salt with the canposition ( Na NO 40% , Na N02 7%, and KNO 53%) • This eutectic cooling medium is heated from 578 oK ( 305"C ) to 598 oK ( 322 °c ) by the recirculation in the shell side of the tubular reöctor. In order to calculate the full dirnensions of the reactor and the nurnber of tubbS required for the reactor we apply the following heat transfer equation

Q

=

U A

.6T

m ( 8 )

U is the overall heat transfer coefficient calculated fr~ the

tube

side and the shell side heat transfer coefficients :

o

33 0.14 . P r · • (}Jo / )lw ) ( 9 ) 1 / Uo = 1 / ho + 1 / hOd + do . ln ( do / d. )

---~----

t dol d . . l/h' d + do/do • l/h. J. J. J. J. ( 10 ) Q = 8, 813.75 KW to be rernoved U

=

120 w/ m 2 °C.

..

- 24

-Hence

Eqn. ( 8 ) gives the total heat transfer area as A

=

4197.024 m 2 .•

Ncw we apply selected tubes of 46 / 50 mn dia. in single pass of and 3.5 m length, then the area of one tube is as follows :

1.25 square pitch

Area of one tube

=

50 . 10-3 • 3.5 •

A

No. of tubes

=

A / area of one tube

=

7631

Then the tube cross-sectional area

= 7\/

4

=

0.55

m

2

The total flow area

=

no. of tubes tube cross-sectional area Mass velocity of the gas mixture

=

44.3418 / 12.682

=

3.49646 sec -1

2

=

12.682 m

-3

The gas mixture density at reactor temperature

=

2.5 to 2.6 kg m

Hence the linear velocity of the gas mixture along the reactor tubes is as follows

Hence:

v

x

=

1.40 In sec -1

~

=

2.5 sec. in the reactor tubes.

-t; r

L

=

3.5 m just right.

For the reactor we apply the bismuth-rrolybdenum catalyst of 0.4 porousity and grains of d

=

3 mn packed in the reactor tubes. Sc for the pressure drop across the reactor

p fran the infonration gathered, we can apply the following equation

d

=

0.003

m

p

L

=

3.5

m

Ap/L

&

=

0.4 1. 75 .

fv .

V

x

2• ( 1 -

~

) = --1 v

=

1.4 m sec x

then Ap / L

=

0.26796 bar or ~ P

=

0.94 bar less than one bar

OK!

0.4531 Bundie dimensions

=

~

=

50 . ( 7631 / 0.215 )

= 5763 mrn

Since the flow through the tubes is single pass, the tubes must he attached to the fixed tube sheets at both ends. With this selected tube dimensions the pressure drop as calculated is less than one bar, care should also he taken in loading so that the pressure drop across the parallel tubes will be equal. This can he achieved by chargi.'1g the same mass of catal vst to each tube, allowing no more than ± 5 to ± 10 % deviation fran the mean. catalyst~ in the tubes rna.y he held in place by perforated retainer plates covered with screen and removable in sections or screen tied to removable grid supports. The grids or retainers are held by attachments to the tube sheet. catalyst can also he held by individual perforated plugs in each tube •

- 25

-The catalyst is durnped by rernoving the screen-support canbination in sections and blasting with air or nitrogen when required. Adequate facilities for inerts and steam required for start-up should be designed as weIl as necessary preheating equipnent. The reactor is preheated by recirculation of rrolten salt bath which is operating near the incipent reaction temperature . The first portion of this reactor serves merely as a preheater, and since the feed and the cooling medium ( rrolten salt operate at the same temperature, this unit at steady state operates for maximum

26 -. '1>ropylene --I~-J' __ I J and etc. I 0_v CAo

•

Sal t outlet (al

GCM'

0 (bI Figure IX.!. C( = X-sectional area

~)

_G

CA/x+dX

0

..

..

..

~ 14--X X + dX L X

The models of the concentration changes in a plug-flow reactor.

Explosion disc

Explosioll disc

- - - -- -

27

-(d) Temperature and conversion profile in the tubular reactor :

An outline of the temperature gradient in the tubular reactor is given in figure (IX.3). In order to calculate the reaction behaviour not only the molar balance in the catalyst bed is needed but also the heat balance of the heat exchanger part of the reactor has to be taken into account. The catalytic propylene oxidation is highly exothermic, first order and an irre-versible reaction so we can have :

( 12 ) RT

Where

-3 -1

- rA

=

Specitïc conversion rate of A ( propylene or else ) mol m sec

Ko

=

Reaction rate constant; sec -1

EAIR Activation energy / universal gas constant. T

=

Absolute temperature

CA

=

Concentration of A mole m -3

The reactor contains 7631 tubes with radius R = 50 mm and length of 3.5 m. The heat exchanging capacity of the tubes is as follows

1

=

--~---U. A ( 13 )

o

.

f

.

C v P H T U 120 . 4197.024 = -332.3435 1560.0 0.97143

While ~ is thought independent of temperature and mixture composition. In this equation ( 13 ), A is the total heat exchanging area of,the all reactor tubes, 7631. The residence time 1r of the tubular reactor ( tube-side ) is 2.5 seconds and the adiabatic temperature rise is as follows :

( - H ) •

c.

fl

T~d =

---fv--~--;~!!---Now we consider the properties of the key component, propylene, in the above equation :

( 340 ) . 3. 2077

--- = 129.41 OK 3.371 2.5

Further, from part (a ) of

cr.

i

S

section, we have seen that the molar balance of " A " over a slice Fdx of th.e reactor resul ts in :

-~-~

(-r A )

t'C

--- ( 15 ) dZ C r Ao

d~

_{( 16 )} dZ