The production of high carbon monoxide

(1)

&

l

/

K'T"

FVO 3103

onderwerp

The production of high purity carbon monoxide

keywords

carbon monoxide, methanol,

pyrolys~ ~ J I. _ t . ~

P

high purity

~~ -ft ~~4.

• ontwerpers

M.W. Kooistra

H-J. Meester

B.F. Tuinstra

Datum opdracht:

Datum verslag:

tll~eJ

tbJ

f-.--~~k

.

015-569854

015-144332

079-213541

12 April 1994

1 November 1994

(2)

1

2.1 Strëams and components sheet part 1

2-2.2 Streams and components sheet part 2

3.1 Heat and mass balance part 1

Lt

3.2 Heat and pass balance part 2 3.3 Heat and mass balance part 3

4.1 Specification sheet of heat exchanger Hl

1-4.2 Specification sheet of heat exchanger H2 4.3 Specification sheet of heat exchanger H3 4.4 Specification sheet of heat exchanger H4 4.5 Specification sheet of heat exchanger H5 4.6 Specification sheet of heat exchanger H6 4.7 Specification sheet of heat exchanger H7 4.8 Specification sheet of heat exchanger H8 4.9 Specification sheet of heat exchanger H9 4.10 Specification sheet of heat exchanger H10 4.11 Specification sheet of heat exchanger H11 4.12 Specification sheet of reactor Rl

4.13 Specification sheet of reactor R2 4.14 Specification sheet of column TI 4.15 Specification sheet of column T2 4.16 Specification sheet of column T3 4.17 Specification sheet of pump PI 4.18 Specification sheet of pump P2 4.19 Specification sheet of compressor Cl

5 Fire and explosion index calculation sheet

2h

6 Hazard and operability study

11-7.1 "Chemiekaart" for carbon monoxide v

7.2 "Chemiekaart" for methyl formate "

7.3 "Chemiekaart" for methanol ".

(4)

1. Introduction

Carbon monoxide is an important raw material in many chemical processes. It is used for the production of phosgene, methanol. acetic acid and aCrylic acid, in the

hydroformylation of olefins to form aldehydes and alcohois and m Flscher-Tropsch synthesis, to name just a few. Usually carbon monoxide is produced by steam

reforming or partial oxidation of natural gas, oil or coal. These prócesses all produce a mixture of carbon monoxide andhydrogeri. callçd synthesis gas. Althóugh in many processes both hydrogen and carbon monoxide are neëded and the synthesis gas is applied direc~y (as in the methanol synthesis process), there are processes in which hydrogeu is

au

unwanted component (most notably in the manufacture of acetic acid). Thus in existing commerciar applications the production of carbon monoxide is in fact separation of carbon monoxide from a synthesis gas stream.

The various separation techniques currently in use are mo&enic separ,,!tion, selective absorption/adsorption, and membrane separation. Cryogenic separation is the

technique used for large volumes and stabie operation. It is also the most energy consuming technique. The two major absorption/adsorption processes, i.e. Eressure swing ads0jttion and COSORBTM, use a copper corppound to selectiv~ly bmd carbon monoxide. oth are most COrDmO!1ly employed in ~ ::ç>~*,oc~~es andlor low concent:Î'ation carbon monoxide streams. Operation IS conSl ere to e ess reliabte' ~genië separàü5'Ïl Dut energy requirements are moderate. Membrane

technology is considered to be the choice of the future, today however the technique is limited to small volume processes, because of the large surfaces of membrane still needed and the high investment requirements. _

,

Our aim is to design a process which produces high volumes of commercial high punty carbon monoxid~ ~) at low energy consumption and reliable and stabie

operation. · ( ~~

(5)

-I

i

2. Starting POi

+

ts for the design

2.1 Product ion of aceJic acid from methanol and carbon monoxide

One of the major high rolurne, commercial high purity carqon monoxide consuming processes is the produq;tion of acetic acid, a chemical of fast growing economical importance. In the U.S.A. 80 perçent of the current operating capacity of 1538 metric mJs oased on meth!mol carbonylation, and an increase in percèntage and capac1ty IS

expected in the near future. Vsually, both methanol and carbon monoxide are deriyed fromh dro en rieh nthesis as feedstock. Because of me necessity of~a hydrogen

ee enVironment 1 cetlc aCl rma on reactor, direct contact of methanol with synthesis gas

i

mpossibl and separation of carbOn monoxide needed . ' , '

Conventionally, 0

e'

echniques aie instalied, but research is'heing done to find

less energy intensive altematives.

-A typical acetic acid plant like the Monsanto production facility in Texas has a capacity of 180.000 tonnely and operates at a temperature of 180°C and 30 bar.

2.2 The MGC pure carbon monoxide process /

(b

è-d)

In the COSORBTM process carbon monoxide reacts wit}{ a copper complex dissolved in an organic phase. Because the activecompontfnt is luted and in another phase, considerate mass transfer limitation can be expecte ,which makes the technique unsuitable for high volume processes. Limitatio ' an he eliminated by using a

concentrated absorbent that i.s in the same' aggt; gation state as the target molècule. In

the 1920's and 60's research was done on re tions of Cl building blocks. Interesting reactions are those forrning h dro en and eth 1 formate from methanol, and

forrning carbon monoxide and met anol om met y ormate; Based on these two reactions, the Mitsubishi Gas Chemical ompany designed

a

process in which hY8Io,[en and carbQ!! monoxide are produced ar seE,arate places, using methanol as

feedstock:

-8rH298

=

52.7 kJ/mol (A)

CH300CH > CH30H + CO 8rH298

=

38.5 kJ/mol (B)

~

\

At one place methanol is converted P methyl formate and hydrogen (reaction A), and at another place methyl formate is

&~~mposed

into carbon mono#de and meth,anol (reaction B). The methanol formed in

iIiê

sec6hd reaction is recycled to the fust. The process became possible due to the development of a catalyst for'reaction B which is still active in the presence of a considerable amount of meth1Ulol. ' , '

2.3 Selection of the catalysts

Goals

As the economie feasibility of the process is important, much effort was put into the

search ~or suitable catalys.ts.ln order to keep the pr?~uced carbon monoxide: as pure

Q

as poss1ble, more emphas1s was put upon the setect1v1ty of thecatalyst than ltS

activity. In fact, a catalyst with a selectivitY-QfJes..sJbM O~ wOQld probably make (

thë plant uneconomical due to loss of m'ëfhanol.

é:::

~

en

VUt

dl .

Literature search stratees ::;>

~

•

The catalysts used in the MGC process are both patented, and therefore little is known

t-of their composition and performance[l]. In an attempt to find'relevant data in the

open literature, an off:·line s~arch Was conducted inShemical Abstracts, To reduce the amount o(,work tgJlJ1l!m...êl~I, only thè abstr~cfs ot 1982-1986 were

considered. The chenllcru'SübSfance'lndex was'searc;hed on the dehydrdgenation of methanol, the produ~tion (manufacture) ofmethylformcite and the pro'duction

(6)

---~I

~~!

!II~~ ~

The production of high-purity carbon monoxide

(formation) of carbon monoxide.;. lso, the author index was searched for relevant articles and patents by MOC.

More recent literature was found using the Science Citation Index. All references to Ikarashi[1] from 1980 until pril 1994 were checked, as weU as all references to

Ushikubo[lO] from 1984 April 1994. Finally, one of the articles [9] was found by a

random library search.

Selection of the deh enation catal st

In the original MO process as described by Ikarashi[1], the dehY~tiOn catalyst

contains c , conium, z~d possibly alw;WWum. A yiela and a

selectivity 90% e reporte<r. .

The litera . ests that copper is the active species in the catalytic

dehydrog~nation of methanol to methyl forma te. A number of carriers is suggested,

such as ~a [3,4,5,7,8] and actiye carbon 18]. The activities and selectivities are \ J;

highly dependent on the compos[tion or the carrier material. Selectivity seems to ~ ~ ~ ~

decrease with increasi~.gacidity, while ~~ inçreases, alth'óügTiihere seems to be ~

some controversy over this [5, 8]. Ion-exc anged CulSi02 catalysts are found to have

a better activity and stability then those prepared by fluid immersIon [3,7]. The most

promising catalyst howeve n' of a..cu 2+ ion exclianged form of fluor tetta

silisic mica (Cu-T$M) ,9]. A s èctivltyof

100 %

is reportëa at a temperature of

~4ö

Ot

and a conversi of 45 %, w ich is better thàn the MOC catalyst. These

figures were measured in' expe' ent where the catalyst was exposed to a 1:2

mixture of methanol in nitrogen gas with a rate of about

~

~~

c

~WX~

second.

11 '"

~

J.JJI

?

,Fn

..

C

!f'

l-

.

Selection of the decarbonylation catalyst

Alkaline salts have long been thought to be active in the catalytic decomposition of

~thYl.i~ to carbon onoxide and met anol [11], and were tested on various

support matenals. sh' ubo [10] founa so Id gO to be very active and 100%

§electixe towards the decomposition to methanol and carbon monoxide. Later papers

on thls subject [12, 14] however, report only a very low activity or ilo activity at all

for undoped Mg~ ___

Ma et al [13] nd that. zeolite exhibited a constant selectivity of 99.5%

at a conver . 0 of ~n temperatuie of 250°C. Because of its high availability

and stabili I , is iStiiëcat yst of choice.

~

.

(!d-

<J~'

~ ,

S'~c

t

hl F -'> Irt ct f)fr/ +

e.o

c::::a ~.rsX )Çöt>é

11To

9f,~~(~t()

C~

J

(Kl.

JIn"a{-

'>

nr-

l'

.U.!

..

cM·~t'l.

.tto°c!

4 -f&ot>(~~

(,4

)

~

(7)

3-3. Thermodynamic properties

3.1 Methanol

Methanol (H3COH, molar weight 32.042 g/mol) is a colourless liquid at ambient conditions.

PL = 791 kg m-3 (at 293.15 K)

lts vàpour pressure can he given by an Antoine equation:

B

In psat

=

A - T +C

for methanol, the following values for the constants can he found: A = 18.587

B = 3626.55 C =-34.29

(3.1)

To obtain the vapour pressure in bars, the result has to he multiplied by 133.32 (bar/rnm Hg).

=

' ~

The liquid viscosity can he correlated by:

The values of the constants for methanol are A = 555.3

B =260.64

3.2 Methyl formate

(3.2)

The liquid viscosity of methyl fonnate (H3COOCH, molar weight 60.052 g/mol) is at room temperature

PL = 974 kg m-3 (at 293.15 K)

The vapour pressure of methyl fonnate can he correlated by the Antoine equation (eq. 3.1) with the following constants:

A= 16.5104 \

~ ~

?

B =2590.87 )

C= -42.60

The liquid viscosity can

Pe

obtained from eq. 3.2 with A = 363.19

B = 212.70

3.3 Carbon monoxide and hydrogen

Carbon monoxide is a gaseous speCies (normal boiling point: ~K), molar weight 28.010. Hydrogen has anormal boiling point 20.35 K and a moraIWeight 2.016 g/mol. . .

(8)

4. Process structure

Fresh methanol is fed to the process t 293K and 2~6 bar. Methyl formate and

hydrogen are formed in the dehydrog ation react r R2, hic~ operates at a

temperature of 503 K and a B e of ar. Car onoxIde and methanol are

formed in the pyrolysis reac r Rl at a temperature of 523 K and a pressure of 19

bar. The methanol formed is ed to the dehydrogenauon reactor. The gasëOOs

reactor products are cooled to condense unreacted methapo nd methyl formate. In

theory separation should he limited to simpie' gas-liquid sep ation: but in pract e

some methyl formate remains in the vapour phase, making xtra effort necess to

obtain the specified purity. The feed stream is split in two d used as'washin fluid

in two meth I formate abso tion column T3:The rodiIct stream fro the

dehydrogenauon reactor IS t9 a stillation column. The top stream consis ng of

methyl formate and hydrogen leaves the column at 395 K dis further cool before

entering absorption column Tl. The bottom stream consis g of methanol d a little

methyl formate, leaves at a temperature -of 434 K and is r ycled to reactor . Heat

is exchanged between the reactor' products and the reacto fee. e reacto -product

stream contains light gasses, which make heat exchange ery har . Therefi e, a heat

pump is used to force heat transf~r hetween the feed an produc streams reactor

R2.

~~

I>--

t<Y/~~

v.(

/f)

~

( h "

~~~

(9)

5-#~f

(

5.1 Heat exchangers / '

~

5. Design of process equipment

Introduction /

Since the reactions in the process occur in the gas phase, while the~ parations require

the reactants to he in the liquid phase, most of the heat exchan~rS are either

cóndènsers or evaporators, or both. In order to con serve ener

y,

the evaporation of the

reactor teêd streams and the (partial) condensation of the actor effluent streams are coupled. An analysis according to the pinch theory sho éd that due to the fact that

gasses are formed in the reactor, it is i 1 letely exchange the latent

heat between reactor feed and efftue (see appendix 8). or the dehy .

reactor (R2) a heat pump was installe solv' Iem.

Since the boiling interval for the methanoVmethyl formate system s very narrow, e

design of the heat exchangers is based on uations for pure com unds:

Heat exchan r Hl: eactor feed de-subc oler and effluent condenser

As there are differences between the hot and the cold stream compositions,

the chosen configuration is a horizontal shell-side condenser. The shell-side

condensation coefficient for this configuration (condensing vapour) on the outside of

one tube is given by [16]: '

(5.1)

The tuhe-side coefficient can he estimated by ~ Nusselt-correlation:

( ) 0,14 Nu = 0.023 Reo.s Pr°.33

:w

h·d· W Cp~ Nu=-'-' Re= Pr=--kf 1t N di~ kf (5.2)

which is only valid in the turbulent flow re~ime. Since the fiuid in the tubes is not

very viscous, the viscosity factor was neglectecL

The overall heat transfer coefficierit can be calculated from:

(5.3)

(5.4)

The mean temperature difference is the corrected log mean temperature difference:

(10)

where (5.6) and JR2 + 1 In (

l--

is )

Ft

=

----~---'----'---2 -

s(

R + 1 -

J

R2 + 1 ) 2-S(R+ 1 +JR2

+

1)

(R- 1) In (5.7) in which (5.8)

Since the tube side heat transfer coefficient is a function of the Reynolds number, which is in turn a function of the number of tubes, an iterative solution method was employed. The number of tubes was estimated, so the Reynolds number and the overall heat transf~r coefficient could be calcul~ted. Thè total heat exchan~llZ area was calculated using equation 5.4. The new estimate for the number of tubes was obtained by division of this area by the area of a single tube. This procedure was repeated until the difference between two consecutive estimates was smaller than one tube.

For heat exchanger Hl this resulted in a design with 1028 tubes, of 10 meters length with a diameter or1(J mm and a wall thickness of 2 mmo In a two tube side passes, one shell side pass setup, this would result in a bundle diameter of 0.63 m and a length of 5 m (exc1uding the heads). The bundle diameter was obtained from the empirical equation (for two tube passes):

(5.9)

The tube side pressure drop was estimated using

<1P,

~

(8.ie

~

+

2.5r~~

(5.10)

where jris the friction factor which has a value of 4.5.10-3 [16]. The pressure drop was found to be approximately 0.28 bar, which is reasonable for a high pressure system.

Since there is condensation in the shell, the shell side pressure drop is much harder to predict with sensible accuracy. A value of 0.1 bar is used as a conservative estimate.

(11)

-7-Since neither stream is expected to be very corrosive and the temperatures are not

very high (well ow 750 K), the heat exchanger can be constructed of plain steel.

(5.11)

This equation is valid for nucleate boiling. If the gas flux at the tube surface gets too

large, a gas film will fonn around the tubes, which will act as a heat transfer resistance. The critical heat flux, at which this will occur is given by

(5.12) It goes without saying that we need to keep the heat flux wen below this value.

For the inside of the tubes, equation 5.2 can be used, if the factor 0.023 (for

non-viscous liquids) is replaced by 0.021 (for gases).

The overall heat transfer coefficient, the logarithmic mean temperature difference and the mean temperature difference were calculated as above. Since the tube side is (almost) isothermal, the correction factor for cross-flow (deviation from true counter

current operation) will be effectively 1.

As with heat exchanger Hl, the calculation of the number of tubes is an iterative procedure. An extra complication is caused by the fact that in equation 5.11 the wan temperature and saturation pressure are needed. Since these are affected by the total heat flux and the heat exchanging area, another iteration loop is needed. Here, the wan temperature is calculated by

T - T

+

U (Ttubeside - Tshellside)

w - shell side h.

1

(5.13)

The saturation pressure at the wall temperature of the mixture was then calculated using AS PEN. The new wan temperature and pressure were entered into equation 5.11, which led to a new overall heat transfer coefficient, and the procedure was

repeated until the change in consecutive estimates was not more than 1 K.

The resulting design is a heat exchanger with 320 pipes with a diameter of 25 mm, a wall thickness of 2.5 mm and a totallength of 10 m. In a two tube side passes, one shell side pass layout, the bundle diameter will be about 0.7 m (Eq. 5.9). The tube

side pressure drop is estimated to be 0.11 bar (using Eq 5.10 with a friction factor of

(12)

o

~

which brin s its tem erature t . Then it is condensed in heat exchanger to evaporate the feed stream. e amount of steam used is fixed by the amount of heat that needs to be transferred in both exchangers. Therefore, not all of the feed stream is evaporated in H3. This is not a problem however, because the presence of liquid in H4 will not influence its heat transfer coefficient due to the fact that the heat flux is supercritical.

The heat transfer coefficients in heat exchanger H3 can be calculated using the equations for pool boiling (Eq. 5.11) and condensation inside horizontal tubes. The flow pattem in the tubes is assumed to be either stratified flow or annular flow. The heat transfer coefficient is the higher value of

(5.14)

for stratified flow and

(

1 +

ffi)

(h) = 0021 kL _ReO

.8 Pr°.43

V

Pv

c . • ~ 2 (5.15)

for annular flow.

The heat transfer coefficients in heat exchanger H4 can be calculated using the Nusselt equations for one-phase heat transfer (Eq. 5.2). Instead, the heat transfer coefficients were estimated and no detailed design was made.

Heat

exchan~er

H5: Pyrolysis reactor effluent subcooler

t

,,~r

Af ter the feed-effluent heat exchangers, the pyrolysis reactor effluent needs to be

(\r

cooled further before it enters the carbon monoxide tower. For this purpose a stream ~ '\ •

of cooling water is used. The coolin water is a med to be available at 2

oe

and 3 bar, and can be heated to 40

0ç.

The process fluid enters the exc anger at 0 bar, so for easy shell construction, it is allocated to the tube side. The tota! amount of heat to be transferred is 8.59 MW, so the amount of cooling water needed is 102.9 kg/s. The shell-side heat transfer coefficient can be estimated using a correlation specific for water:

h_i= 4.2x106 ( -4.11

+

0.02 T) U~·8 di 0.2 (5.16)

The tube-side is condensing vapour, so we use the higher value of equation 5.14 and 5.15. This leads to a design with 500 tubes of 10 m length with a diameter of 25 mm and a wall thickness of 2.5 mmo The tube bundie diameter is 1.15 m, and the tube side pressure drop 0.0035 bar, which is extremely low.

Heat exchan&er H6: Dehydrogenation reactor feed preheater in heat pump cycle The hot recycle stream is mixed with the cold bottom stream from the carbon monoxide tower just before heat exchanger H6. Due to this fact the amount of heat required to bring the stream to its boiling point is small in comparison with the total amount of heat transferred, and the equation for pool boiling can be used for the entire heat exchanger (Eq. 5.11). The fluid in the tubes is condensing steam, so we use Eq. 5.14 and 5.15.

An iterative procedure leads to the following design: 465 tubes, 7.5 meters in length, with a diameter of 50 mm and a wall thickness of 2.5 mmo The bundle diameter is 1.6 meter, and the tube side pressure drop 0.04 bar. The fluid velocity in the tubes should not exceed 10

mis

in high-pressure heat exchangers, hence the larger tube diameter. Heat exchanger H7: Dehydrogenation reactor feed preheater

(13)

-9-As the amount of heat that is transferred over the heat pump is fixed by the amount of heat that is released by bringing the effluent stream to its target temperature, an additional heat exchanger is needed to bring the feedstream to its target. As in heat exchanger H4, high pressure steam is used. Aside from a slight difference in stream composition, the operating conditions of heat exchanger H7 and H4 are identical, so the same heat transfer coefficient can be used.

Heat exchanger H8: Dehydrogenation reactor feed cooler in heat pump cycle In this heat exchanger the reactor effluent is cooled to its target temperature, while water at 1 bar is evaporated. For easy construction of the heat exchanger shell, the Iow pressure, Iow temperature water is allocated to the shell side, and the hot reactor effluent to the tubes.

The dehydrogenation reactor effluent enters this heat exchanger as a superheated gas. The heat capacity of the gas is Iow, so the amount of heat that needs to be removed in desuperheating is relatively small compared to the amount of heat that is liberated in the condensation of the gas. Furthermore, the gas will eondense onto a eoid surfaee, so that surfaee will be completely covered by a film of the liquid at its boiling point. The result of this is that the heat transfer eoefficient can be estimated using the equations for condensation inside horizontal tubes (Eq. 5.14 and 5.15). The shell side heat transfer coefficient can be estimated using the equations for pool boiling (Eq. 5.11).

For a two tube side passes, one shell side pass design, the dimensions are as follows: 725 tubes, with a length of 10 m, a diameter of 25 mm and a wall thickness of 2,5

mmo The shell diameter is estimated to he exactly one meter, and the pressure drop will be about 0.6 bar.

Heat exehanger H9 and HIO: Column 1'3 reboiler and condenser

No detailed design was made for heat exchangers H9 and HW. The heat transfer coefficients where estimated from literature data [16 fig. 12.1].

Heat exchanger H11: Column TI top stream subcooler

Heat exchanger H9 can be designed in the same way as heat exchanger HS. This leads to a heat exchanger with 700 tubes of 10 m length with an inside diameter of 25 mm

and a wall thickness of 2.5 mmo For a two tube passes design, the tube bundle diameter would be 1.3 mand the pressure drop 0.015 bar.

5.2 Reactors

--

<-~"Y~

'!

Introduction

Since the avaiIable data is scarce, the re§lçtots can not

bs

desigued in detail. HardIy any data at all is ~vailable on the kinetic rat~ of the reactions, so lh~~U~ of cat t is kno n and the reactor vo u e can not calcu at .

Here an attempt is ma e to m e a rougli estimate 0 ereactor' mensions based on

literature experiments and heat transfer requirements. It needs to be stressed though, that the reactor dimensions are based on,gench-scale laborat07u ex~erinients, so due to scale-up effects, they couid easily be a factor 2 or

4

wrong usu Iy too smali).

- -~

The pyrolysis reactor CR 1)

~s Lee at. al. [12] ~eport, the equilib~um for the decomEosition of ~ethyl formate mto carbon monoxIde and methanol hes complete1y to t e product slde for pressures up to

50

bar at 500 K. The testing of the catalyst was done by Ma et.' al. [13, 14] on a bench SëaIe floWSy'Stem. A stream of 0.-5 ml/s of a 26% ' eth 1 formatè-ar on mixture was led over 0.2iof catalfist at

523

K and áhn sheric ressure. s amounts to a S e 6 5 I

mI

met yI formate per seco per g cat . The

conversio~ wa: 17%, at is 110)n1/(s kgçat). Assuming the s rate can btained in a commerci or (at the prbcess pnessure!), the tof càtalyst requ ed to convert a~ut 0.1 kmol meth 1 formate er second is a t 2 tons. Assuming ,

(14)

The production of

high~purity

carlx," monoxide

~~

catalyst bed density

O~fOOO

kg/m3 and applying a safety factor of 2, the r

é

volume would be abo t 4 m3. r o r

The decomposition re . s end?the~c, so for ~sothermal reactor erration, heat

needs to be transferred to the reactlOn nnxture. The amount ot heat.ne ed IS

~:Jt

MW·

The heat will be transferred using Dowtherm heat exchanging fluid.

S

IS a mixture of difenyl and difenyloxide that condenses at 587 K at 3 bar.

The total heat transfer coefficient can be estimated to be about 550 W/m2 K [16], and,

as the corrected logarithmic mean temperature difference is abotit 64 K, a heat

exchanging àrea of about 100 m2 will be needecl. For a tube-and-shell type reactor,

the tube diameter would

Së

about 150 mm, and the total tube length 225 m (i.e. 45

tubes of 5 m length'. . ' . . . .

The Dowtherm could be (re-)heated in a furnace, using 464 kg/h fueloil (with a

fumace efficiency of 70%).

t.nrf?

The dehydrogenation reactor (

~

'2.

')

The dehydrogenation is an eqcinibrium reac ·n. At a temperature of 503 K, about

5i%

of a pur~ methanol stream can be con erted to methyl fonnate. 'As"tforikawa[9]

s ows, aböut ~ conve . - s a . n a laboratory setup where 700 mg of the

catalyst was exposed t 1 ml/s of a 1:2 thanoV' 0 e . a s ' at 13 K.

This means that a 00 tril methanol was con ert per second per kg catal yst. 0

convert 200 moVs a ut 20p0 kg catalyst would --lleed~. Assuming a safety factor

of 2, and a bed ns! of 1, the· total rèactor v me is abou~ m3;--. --_ .

As hot utility fl ld, hig - ressure steam 4 , 683 K) can used, since it · - - .

----condenses at 5 3 K - 20 eactor temperature. If we use 6.88 kg/s of HP

steam, cool it t, 523 K and condense about 25%, the heat requirement of the reactor @ê)

are satisfied. The condensed steam can be used as heating fluid for heat exchanger 7 so the reactor and its pre-heater will be effectively integrated.

The energy needed for isothermal operation is 6.1 MW, so - assuming an overall heat

transfer coefficient of 250 W/m2K - the total heat exchanging area needed is 330 m2.

This leads to a reactor with 433 tubes, 5 meters in length with a diameter of 2.8 cm. In

order to keep the pressure drop within bounds, a better setup would he 188 tubes with

a diameter of 100 mm and a length of 5 m. This would take the bed volume up to

about 8.25 m 3.

5.3 Separation equipment

In conventional techniques high purity carbon monoxide is made by separating

c~n m9nQxi~e from ~ngas' which requires cryogenic distillation.-As mçntioned

he ore cryogemcs.are energy consuming, so this process lias bëen designed·to

produce hydrogen and carbon monoxide at separate locations in the plant. Therefore <::::J

in theory the com onent se aratio h re ce to . a -li uid se tion. ~

Unfortunately a tt e me yl ormate remruns in e gas p ase, ng extra e fort

~:~~ary to maintain the total efficiency of the pr~ess and to ~\~~\)e high purity

Methyl formate absorbers in product streams

\)~

I

\}'lN

,

To remove the remaining methyl formate from the product stream one CaIl decide to

lower the temperature in order to shift the equilibrium towards the liquid phase.

Lowering the temperature 'means also grè~ter temperature differences in the prqcess

loop and an equilibrium shift of hydrogen towards the liquid phasè. Therefore it was

f

decided to ash the r . streams with the feed str ams. By adding fI:esh liquid ~

removing saturated 1t is pO~SI e to 0 tam gas . ase me yl fonnate conçentrations

lower then the equilibrium values.. at the same onditions. . ' •

The column diameter was calculated by divid g the vapout flow rate by the maximum vapour flux. The maximum vapour ux is givt?n by:

(15)

-y.

=

w

K4Pv(PL

-pJ

42.9 Fpo.1L / pdO.1 (5.17)

where F_pis a packing factor which is a characteristic of the size and the type of

packing, K4 a constant dependent on the flooding velocity and the pressure drop per height bed. The densities and viscosity were calculated by Aspen. Designed was for

40 mm water pressure drop per meter and 38 mm metal pall rings. The height of the

1

column was estimated by using Cornell's method [34] as Z

=

10 m, for a number of

stages Nog

=

10 and a height of transfer unit Rog

=

0.9 m. ~~ .

Methyl fonnate methanol distillation column

In order to reduce the flow through the equipment and to improve the operability a distillation column was added to the process sheet (TI). The intention of TI is to maintain the composition of the feed of the decarboxilation section and to recycle methanol to the dehydrogenation section. Estimations of the number of trays, the reflux and reboil ratio, the top and bottom stage temperatures, and the flows were made with the aid of Aspen. Between 0 and 30 stages an increase in the number of stages resulted in an improved separation, above 30 stages hardly any influence was experienced. Considering the separation the number of stages should be at least 30. The more stages the lower the reflux ratio, resulting in an decreasing energy

consumption. It was decided to take the minimum amount of stages because the molar reflux ratio of 4.3 was still very reasonable.

To calculate the column diameter an estimate of the cross-sectioned area is needed. The net area can is calculated by division of the gas flowrate by the maximum vapour velocity. The maximum vapour velocity can be taken as 85% of the flooding velocity which is obtained from:

(5.18) Kl can be obtained from Fig. 11.27 [33]. The liquid-vapour flow factor FLV in Fig.

11.27 is given by:

(5.19) As a fust estimate a tray spacing of 0.5 m was taken. The flooding velocity was calculated from eq. 11.81:

base Ui = 1.07 ms-l top

Ui

= 2.36 ms-l

When 85% flooding is considered to be reasonable, the vapour velocities become:

base Uv = 0.85 . 1.07=0.91 ms-l

top Uv = 0.85 .2.36=2.01 ms-l

The volumetrie vapour flow rates were calculated by Aspen: base Y v

=

4.11 m3s-1

top Y v

=

5.43 m3s-1

By dividing the vapour flow rates by the maximum vapour velocities, the net area's of the distillation column were calculated as:

(16)

top An = 5.43/2.01 = 2.7 m2 bot An = 4.11/0.91 = 4.5 m2

When 12% of the column area is used as down corner area, the cross-section area becomes:

top Ac

=

2.7/(1-0.12)

=

3.1 m2 botAc = 4.5/(1-0.12) = 5.1 m2

The column diameter was calculated from the cross-section area: top

de

=

«3.1 . 4)/p) 0.5

=

2.0 m

bot

de

=

«5.1 . 4)/p) 0.5

=

2.5 m

The same diameter may he used above and below the feed, reducing the perforated area for plates above the feed.

The hole area is determined by the operating velocity and the weep point. For stabie operation the hole area should be designed to give a weep point which is weU helow the minimum operating velocity. A hole area of 5% of the active area satisfies these conditions.

The active area is given by:

Aa

=

Ac-

2Ad

=

(1-0.24)

Ac

(5.20) for a column area of 5.1 m this results in an active area of 3.88 mand a hole area of 0.194.

The weir length was taken from Fig.11.31 [16] as 1.9 m. Estimates were made for the weir height: hw = 50 mm, the hole diameter: dh = 5 mm, and the plate thickness is 5mm. The numher of holes for these estimates is Nh

=

9.8 . 103. The pressure drop was calculated to he 2.78 . 10-2 bar per plate. The plate efficiency was estimated to be 0.75. The total pressure drop will he

LW = 2.78 . 10-2 • 30/0.75 = 1.12 bar. (5.21)

5.4 Pumps and compressors

In the flow sheet, three process cycles appear: the main loop for returning methanol formed in reactor R2, a recycle loop over reactor R2 and a heat pump lOOD. Each

cycle r uires a urn or com r 0 stimulate fl f r ess ft · The energy

W

requrrement or a specl pressure difference and mass flow is much lower for a pump than for a compressor, and whenever possible pumps should he chosen. In the main loop and the recycle loop liquid steams are available and pumps could he added.

1\

compressor was added in the heat pump loop. The theoretical power of a.gump or compressor IS glven 6y: - :

=::5-==

~p

i

, _.

.

I

~

'P

W.=$.pg

!

~

;1 &~

,

(~ .~~~

Ij

The total efficiency was taken as 0.75, giving a shaft power of Ps

=ü

.75 W_s

To prevent cavitation in the pumps a safe difference hetween the inlet pressure and the saturation pressure has to he kept. This available Net Positive Suction Head is defined as:

(17)

13-(5.23) The stream from distillation column T2 leaves the column at its saturation point,

which means that there is no NPSH available. Special care should be taken in the

(18)

6. Balances and equipment specifications

The streams and component sheets as weU as heat and material balanees and equipment specification sheets are given in appendices 2, 3 and 4.

'11

~ -JI

Ol{

--~--l

bo

..

(19)

15-7. Process control

In designing control schemes different objectives should be taken into consideration: -safety

-production rate and quality -economy

Each of these objectives can be implemented at different levels of process con trol. The highest level concerns the overall mass and heat balances. The variables related to the overall con trol are the reaction conditions, the feed rate pressure, temperature and composition, the steam temperatures and pressures, and the product rate,

pressure, temperature and composition. These are related and therefore the designer is limited in his choice of setpoints. Lower levels of con trol deal with the steady and safe operation of process loops and equipment. Variables related to lower levels of control are dependent on variables of higher levels and the setpoints are dictated by the choice of the overalloperation.

The most important setpoints to con trol are the product rate and composition, but these points are fixed by the reaction conditions. When the reactor size and catalyst activity are given, the choice of temperature and pressure fixes the whole process. The overall mass balance can be controlled by measuring the contolled variables, thus establishing the molar rate of (bound)carbon out of the process and compare this rate to the (bound)carbon feed rate.

7.1 Reactor and heater con trol

The vast majority of vessels in the flow sheet are reactors and heat exchangers. The reactors and heat exchangers have the same con trol objectives and therefore could be treated alike. The objective is to raise the temperature of a process stream to a desired level and to keep it at that temperature.

The available controlled variables are: -vessel pressure

-steam pressure

The temperature of the condensating hot stream only depends on the operating

pressure. By controlling the operating pressure one basically controls the temperature. The available manipulated variables are:

-vessel flow rate -condensate flow rate

It is obvious that the arrangement should combine the vessel pressure with the vessel flow rate and the steam pressure with the condensate flow rate.

7.2 Distillation column control

The distillation column was incorporated in the flow sheet to reduce the flow through at least one reactor section and both absorption columns. The objectives in the control of distillation column T3 are to maintain the distillate flow rate and composition at the desired value. According to Stephanopoulos[ 17] four variables may be controlled. Usually two level controls will be added.

The available controlled variables are: -light-end composition

-distillate rate

(20)

-liquid level in bottom section

The available manipulated variables are: -distillate rate

-reflux rate -steam flow rate -bottoms flow rate

For a manipulated variabie the effect of a change on its controlled varia bIe should be as direct as possible. Unfortunately only three pairs are available with direct response,

i.e. controlled variabie and manip~lated variabie at the same side of the column. The

best configuration is the one with the least indirect and the least difficuIt response over the column. No least indirect and least difficult response can be made by manipulating the bottoms flow rate nor the distillate flow rate. This leaves the steam flow rate as the indirect manipulated variabie. The easiest response of the steam flow rate is on the distillate flow rate, making the best couple for indirect control. The easiest way to con trol the distillate composition is by manipulating the distillate flow rate. The accumulator level and the reflux ratio are left to form the last coupie.

7.4 Absorption column con trol

The objective of adding the two absorber columns was to reduce the loss of methyl formate over the top. The control objective would be to keep the vapour composition at a specified value. However the methanol feed rate and the bottoms flow rate are fixed outside the column and therefore only the column's pressure and liquid level are subject to con trol.

A vailable controlled variables: -column pressure

-liquid level at bottom

A vailable manipulated variables: -vapour flow rate out

-liquid flow rate out

It is obvious that the column pressure should be controlled by the vapour flow rate and the liquid level by the liquid flow rate.

7.5 Pump control

The objective in pump and compressor con trol is to keep the outlet pressure at a desired value. Pumps and compressors with adjustable numbers of revolutions are to be preferred above fixed speed types.

A vailable controlled variables -number of revolutions

-A vailable manipulated variables

(21)

8. Safety, health and environment

8.1 Hazards

The compounds present in the process are hydrogen, carbon monoxide, methanol and methyl formate. Copies ofthe 'Chemie kaarten' are provided in appendix 7. Most deviations from normal operation will have direct effects and concern fire and

explosion risks, and short time exposure to high concentrations; those risks are called safety risks. Some risk is involved with low intensity exposure to carbon monoxide duririg moderate time, this will be called a health risk. No very long term low

intensity hazards, which can be addressed as environmental hazards, are present in the process.

8.2 Toxicity

Severe safety hazards are involved with the toxicity of the compounds in the process. MAC values range from 25 ppm for carbon monoxidt to 200 ppm for methanol. The carbon monoxide value is extreme and special instructions should be followed.

8.3 Hazards involving loss of containment

Loss of containment usually occurs as a consequence of heat and/or pressure build up in the system. Under normal operation, hardly any side reactions occur and, as the main reactions present in the process are endothermic, build up hazards will be small.

Deviations from normal operations inside the process are summarized in the Hazard

and Operability study in appendix 6. Hazards occur when oxidizing components enter the process or when mixtures of hydrogen and carbon monoxide are formed. The reaction of hydrogen with carbon monoxide under formation of methane and water is highly exothermic, but usually not explosive, because of kinetie constraints.

However, the influence of the catalyst on this reaction is unknown and care should be taken. General information about the hazards of explosions are estimated by a method

developed by DOW, and is given in appendix 5. A safe estimate of the DOW fire and

explosion risk index was found to be 252.8. Values higher than 90 are considered to be extreme, and for this process this means application of all preventive and

protective measures available. Special relief devices should be instalied to prevent extreme pressure build up. The escaping gasses are very toxie and flammabie, and should be vented to a flare stack.

(22)

~~~~~

9.1. Introduction /

9. Process Economics

~

To be able to compare the cost otF ucmg c -monoxide with acquiring CO on

the market, we need to establis what the productio osts are. These were calculated

using themethod descFibed' Chemical Engmeenng, ol. 6 [18]. This includes a

factorial method for estimatin the investinent needed For comparison the

investments were also estimat sing Iayler' s me , as described in De

Chemische Fabriek, part 3 [19]. -'--~

9.2. Break-do...-.-~~'

The total co st f

costs (Cv) an s

Fixed costs inclu . hl charges, maintenance, royalties, amongst others. Variable

costs are mainly (raw) material costs. Semi-variable costs are sales, genera! overheads

andR&D. ( ' .

9.3. Investments

~

w:ti

&..-~ ~

la. Factorial method

In the factorial method of co st estimation the pxed capital cost of a chemical plant is given as a function of the total purchase equipment co st 6y the equation:

Cfc =Cpe' fL (9.1)

where Cfc is the fixed capita}. cost, Cpe is the total deliyered cost of all the major

equipment item~ (reacti~n vessels, columns, heat exchangers, e.tc.) andfL ts the

so-called Lang factor. The Lang factor has its basis in the following equatlöii:

IL=

(1 + 11 + 12 + 13 + 14+ 15+ 16+ 17+ Is+ 19) 0+ fIo+ in +

flv

(9.2)

direct casts ( Î

h

=

0.4 (equipment.erection) - (

f2

=

0.7 (piping).?- n

,\t.'l::. •

f3

= 0.2 (instrumentation) .. ,.. ~ ~,~ f4

=

0.1 (electrical) ~ ~ " , - ...,. c,'

A

fs

=

0.15 (buildings, process) 9

(.0

1 ~ f6 = 0.5 (utilities) ~ .. .., '\. ./

h

=0.15 (storage)

~

/ ---fs

=

0.05 (site development) ".. / f9 = 0.15 (ancillary buildings) \

4?

.

indirect casts . ,. ./'7 ' .

ho

= 0.3 (design and engineering)

~

0 .

IV' , '

f11 = 0.05 (contractor's fee) ' /' /

f12

=

0.1 (contingency) ~ !

(These numbers are valid for a plant that processes fiuids) /

the equipment, a con enc~adds another 10% to the to.tal cfu t ~t. The Lang

factor therefor IS 4.93. . ~

=

,

1-fJu,~~J.c-·

~

~ -

(23)

19-The total equipment co st for this plant is:

f

1,260,000 Heat exchangers ~ ~

f

42,500 Absorbers c:;::--:::'

f

435,500 Distillation column

-

D

f

190,000 Reactors

f

897,000 Compressor & pumps

---(

zt .

~.,

;

f

2,825,000 Total

U sing the Lang factors ...

W;

arrive at a figure ....

1 -

3 -

.9

-

rm

-

l

-

li

-

o

-

n

-

~

-

or

-

to

-

ta1

-

fi

-

x

-

ed

-

c

--

·

t-al..J.

~

When w.e ~stimate working capital as 10% ''lf nxea, our tot mvestmept adds up to

f

15.32rmllion. ' - - - -

(7-

-

ÇO"

).

~~;

-f

/y~

2. Taylor' s Process Step Scoring ot

k

.

Taylor's method, published in 1977, gives an estimate for the investments 'battery-limits' (64% of total investment [25] ) based on global process steps. Each step is rated according to temperature, prèssure and materials, but also relative to throughput. Taylor's equation is:

Ib = 45f . p0.39 . CI/ 300 IJC ~f 4.- (9.3)

where Pis the plant .capacity in thousands of tonnes per year, Cl is a co st ~ndex

- N

(1977=300) and fis the costliness index, ~efined as the sum

L

(1.3)Si.

# nE 1

-The calculations are given in table 9.1. Table 9.1. Calculation OfIb according to Taylor

Throughput Materialof Reaction/ Pressure/ Other Tota! score

Construction storage time Temperature Storage/ Handling H2 -3 0 0 0 1 -2 (exoiosion) 0 ) ₀ ₀ ₀ ₀ ₁(toxic) 1 Gl40 0 0 0 0 0 0 Me-F 0 0 0 0 0 0 Reactions washH2 3 0 0 0 0 3 react 3 0 0 0 0 3 washCO 3 0 0 0 0 3 react 5.5 0 0 0 0 5.5 column 5.5 0 0 0 0 5.5 Tota! ~1"" CostIiness index 0.591 1.3 1 1 2.197 2.197 2.197 4.233 A.211 18.95

--

-Ib is therefore: 45 . 18.95.92.2750.39 . ÇI/300 = 4979.7 in thousands ofpounds (for

1977). Again the right cost index was not available at press time. The PE composite

plant construction cost for the UK in 1977 was.llQ [26]. For 197~ the same index

was1Q1. We therefore claim that this plant woUlà'have cost .5.8~ million pounds in

1~79':'"Tö go to 1993 guilders we can again use th~ 818% used above, arriving at

f47.7 million.

As this is only the investment 'baU1f.§-limits' we have to multiply by 1/0.64 to get the

total investment, which gives us

f7 .

rmilion. \

-It is apparent that there is a large discrepancy between the more accurate method,

using detailed equipment intormatlon ana Lang factors, and the step me~hod

described above. This is probably due to the 'author's inex erience with Ta lor's

method, and the lack of information on now to mterpret t e p ant accor ng to t e numbêr of process steps.

(24)

Because the detailed method using Lang factors is more credible we will discard the result of T;ylor' s method.

;4.Fixedcosts

(d

17)

~d

-

.

Fixed costs are those that äon't depeI1lton the amount of product produced. These are:

/

Maintenance (labour and materials) , Capital charges

Insurance .

<'

~ Rates (loc al taxes, larid use)

10% of fixed capital see below

1 % of fued capital. ~ il

2% of fixed capital

4~

1 % of ~ed c~pital

Wessel equation -see below

0.10

().IO D. f>q

->v;>

Royalties.and license fees

Labour Catalyst'~ \ f 1,390,000 f 1,390,000

f

139,000 f 278,000

f

139,000 f 700,000 f 192,790 f 4,228,700 se~ belo~. ' '

-o

.2-4

~.

00

Total

The capital charges are taken as 10% of fixed capital. This means that the plant will be written of on a linear basis in ten years. e

t

:

0.30

'2,b

For labour estimates, we used the Wessel-equation [27] :

L = 32 . N . pO.24 _(9.4)

Where L is the labour cost in thousands of Dutch guilders (1986) and Nis the number of process steps. The number of process steps is 5 (two rea,ction steps, gas cleaning, distillation and héàt exchanging / heat transfer) and the plant capacity is 92.275 kT. Operating labour is therefore f475,000 (1986). To estimate the infiation'we usëd the PEP German construction labourprice index' [23]. For July 1993 this is 733.2, and for July 1988 it is 559.9. This give~ us

all

annual increase of

.1,2!%.

Therefore we

multiply the 1986 labour cost figure w i U ' 547) and round up to arrive at an L

estimate for the late 1993 labour costs 00 000. ' .

i>.

-r

Th..e

~

ca®.~rU"s

·

. the plant, we estim Wl e a usefullife

~e

(

SÓdium-13X zeolite's a widely ~vailable catalyst and will proba s ~an

r

15 er kil . e Cu2+ - TSM catalyst"wiH.--pFO bly have to he

1:0

order )~ ~"-an product! n is t straightfo~ard, like impregnating . carrier material. Therefore .. ~

~Q.X.li(...I.<.U,4d1. :)~1 a price.9f '70 er kilogram for this catal. As both reactors take a load of 4 tonne càt, the t t cat ys . 0 IS r f192,700 per year at a f/$

exchange rate of 1. ' ___

.-.;;;;;;c:-The amount of Dowtherm e estimated because there q

available about the volumetric size' of the heating oil system. This depends heavily on ~ _ \

the physical distance to the complex's furnace. In addition to that no price for~~ ..

Dowtherm A was available at press time, so this co st has been left out. ,

L..--?

D~)

9.5. VariabIe costs .

r

.r

jJJ ,

The varlable costs include raw materials u . . . 'sceHaneous materials.

l

The price of methanol is ai a glob igh right now and not very useful to make accurate estimates as far as planning a new plant is concemed. Mid-July 1994 the methanol &pot price was DM 690 per tonne, while third quarter contract prices were _ rumoured to be as high as

DM 1000

..

Historically though, the meth~ol price has hovered around .... $0.20 per kilo~am [28] and in 1991 was as low as 25 29.

Assuming that markeis wIlt setile aown again soon, we Wl use er tonne as Qur

best estimate of future methanol prices.

The hydrogen produced in this process is pure enough to be sold as a product, and therefore gets the fuH by-product credit. We estimate the hydiogen (0 bnng In j1800

per tonne, as this was an accurate price. in 1992 [30]. .

Utilities consumed by this plant will he charged at the appropriate price as listed by the Dutch AssoCÎation of Cost Engineers for November 1992 [31].

(25)

22-The production of high-purity carbon monoxide

The heat supplied to the DowtheIm A heat transfer oil is thought of as supplied by heating oil, with a lower heating value of 41.45 GJ/tonne. The heat required is 3.74MW; ifwe assume a heating efficiency of 70% this amounts to 3712.3 tonnes of heating oil.

An overview of the variabie costs is given in table 9.2. Table 9 2

..

' variabIe costs

Amount Cost ( 106

_{f /}

_year)

Methanol 92.28 . 1()3 to~/y 27.684

Cooling water _{24.15' 10}6

_n:i

3_/y 2.415 I-- _')

Electricity 485.8 kW 0.505 ~

_I

I

HighP steam _{1.053· 1()6}_to~/y 36.86 ~ .., '1{;;,

'1'

Heating oil 3712 tonnes 1.13

"

l

Tota! 68.59

Hydrogen 15.24 . 1()3 to~/y -27.42

Net variabie costs ~

41.17

9.6. Indirect production costs

As this plant was conceived as part of an acetic acid plant complex, sales and research & development are considered to be negligible.

9.7. Product

Cost

~ttt..l.,)

Total production costs are Fixed:

f

4,228,100 _ ____

Variabie:

f

41,170,OOO~ Total:

f

45,398,700

This is for a production of O. 96153 kmol er second, which is equivalent to 8468.6

Nm3jh, or 67.75" 106 _Nm3_/y. _{oduction cost is erefore}_fO.67_{per Nm}_3.

--=-=

9.8. Competing Processes

KTI BVr321 supplied data to compare this process to existing technologyies being C cryogenic sèparation, membrane separauon, pressure swing adsorption, methane wash

and their propriety COSORBTM reversibie complexation technology. These figures are for a smaller plant, 3000 Nm3, for CO at 36 bar in 1987 guilders.

Table 9.3 CO

productiOI

~

hnolo

~

omparison

.

~,...

-COS ORB Cryogenic Methane Membrane PSA

H~~ TM _separation _wash Manufacturing cost 0.670 , 0.696 0.698 0.712 0.700 0.79,8 Cf/Nm3) i"'"

-

-,.,

.

COpurity, % 98.2 97.9 96.8 ( 92.8 ~97.0 96.7 Major impurities CH40,H2 C02 CH4,H2 - . y' "-"-CH4 CH4 ~~4, n:}.

From the table it is apparent that the MGC process designed in this study can compete with present day instalied (proven) technology, both in manufacturing cost and

product purity. However, for the MGC process a write'off time of 10 years was used, but 4 years for the other processes. However, n ' t is' resented here has room for imRrovement. This will be c an e n chapter·1 .

""The Return Un Investment (ROl) was not cal . e was no price available for over-the-fence ÇO. However . we assume a WIJt]

á

%

'

(that is, profit' is 15% of investments) we should m a profit f2,298-,0UU: ThlS would mean our selling price for carbon monoxide s 0.704. N t adjusting' fQr the general costs this se~s our Pay-Out-Time (POT) to 6.7

=....

(26)

-~

!t.o.<\.\ "- LOf.

tJ

W ....

·

9 ~

10. Discussion and ecommendations

/

V"t..

~

The aim of our design wa e production of high volume commercial Igh purity carbon monoxide. In 011 r to be abl~ to compete with a proven tec ology like

cryogenic separation, e process needs to be not only more cost ficient, but also stabie and reliable.

The energy consumption is an important factor in the econ y of the process. In the current process structure, as a result of the low conversi in the reactors, the (eCacle stteams are lar~, and the main stteam is consequentl evaporated ánd condense seveiä1 times. e combination of these two factor . s the cause of the high' energy requirement. In this chapter some changes oséd to improve the process. The conversion in reactor R2 is onl abo e reason of this low conversion is the fact that the dehydrogenation is an eq " reaction. One way to force the

equilibrium to the product side is to reduce the pressure. However, this will cause the

0

stteam volumes to increase significantly, resultmg in large es i ment. Also,

the acetic acid plant requires carbon monoxide at a press e of 30 bar,

sa

larger leed

compressor will he needed. ( -

=

,~

,

~

carbon monoxide

R2

Figure 9.1. Schematic representation of an altemative process structure.

A second way to shift the equilibrium is to reduce the amount of methyl formate in

~

edstteam of reactor

1U.

There are two ways

in

whidi dUs can he accomplished. One s to place the ethanoVmeth 1 formate distillation column TI hetween reactor

g

and reactor R}Jfig. . . -nt IS way, t e unreact m~t y ormate can he recycIea fo the te of react9r Rl while reactÇ>r R2 can he fed with nearly pure methanol. The conversion of methanol in reactor R2 will.increase to neady 50%. reducing the recycle stteam to about 25%.

No reliable data is àvailable on the kinetics of the pyrolysis~reaction. In our calculations we limited the conversion tp 50%. However, if the reaction is fa st enough, the conversion ofreactor Rl can reach 100 %. In this case, no methyl formate can reach reactor R2, and the recycle stteams will decrease by about a factor 7. For this situation, the presented process structure is optimal.

(27)

24-hydrogen carbon monoxide

Rl R2

Figure 9.2. Schematic representation of an altemative process layout without distillation column T3.

In case the conversion in re,actor Rl is lOO%, it could he considered to remove @stillation column T3 altoge,h

_ern

This would increase the recycle streams somewhat

comparêd to ihe last situation (the total decrease is a factor 5.8) but the energy requirements and the investments will decrease significantly. The main energy consumption in this process scheme (aside from the heat of reaction) will he caused by the subsequent evaporation and condensation of the reaction mixtures. H one could find a catalyst that is still active at low enough temperatures, the reactions could take place in the liquid phase. Then, there is no need to evaporate the recycle stream, and the total energy consumption will drop to about 10 MW (the heat of reaction). In cryogenic separation only physical transformations take place. The chemical reactions that are the basis of our process add an extra complexity. This cou1d increase the risk of instability by catalyst deactivation and the build-up by-products. In order to be competitive, the inherent stability of the cryogenic separation must be met by engineered stability in our process.

(28)

11. Conclusion

The production of high volume commercial high purity crabon monoxide from

methanol is technically possible. In this form the process is not economically feasible. However, it is expected that with certain improvements the process can he very competitive.

(29)

26-Nomenclature

Lower case symbols

de

Column diameter di Inside diameter

do

Outside diameter

f Costliness index (Taylor)

fL Composite Lang factor g . Gravitational acceleration

he

Heat transfer coefficient for condensing vapour hi Inside heat transfer coefficient

hnb Heat transfer coefficient for nucleate hoiling ho Outside heat transfer coefficient

had Outside dirt heat transfer coefficient Ib Investment, battery limits

kf Thermal conductivity of a fluid kL Liquid thermal condictivity kw Thermal conductivity of the wall qc Critical heat flux

Ut Tube-side fluid velocity

u_f Flooding velocity

Uv Superficial vapour velocity Upper case symbols

A Heat transfer area Aa Active area of plate

Ac Total column cross-sectional area

An Net area available for vapour-liquid disengagement Ad Downcomer corss-sectional area

Cf Fixed co st

Cfe Fixed capital co st

Cpe Purchase process equipment cost

4

Semi-variable or general co st Ct Total co st of a product Cv Variable co st

Cp Specific heat

F_{LV Column liquid-vapour factor} ~ Packing factor

1't Correction factor for the log mean temperature difference Kl Constant in equation 5.18 K4 Constant in equation 5.17 L Labour cost

Lw

N N P psat ~P

Q

Si

Liquid mass flow rate Number of tubes

Number of process steps (Taylor) Plant capacity

Vapour pressure Pressure drop

Heat transfered in unit time Complexity score (Taylor)

m m m ms-2 Wm-2 K-l Wm-2_K-l Wm-2K-l Wm-2 K-l Wm-2 K-l

f

kg m-1_s-1 Wm-l _K-l Wm-l _K-l Wm-2 m s-l m s-1 m s-l m2 m2 m2 m2 m2

f

kJ kg-1K-l K

kf

kg s-1 kt/yr bar bar Ws-l

(30)

öT m Mean temperature differenee öThnLog mean temperature differenee T_w Wall temperature

'{'Sat Saturation temperature

V v Volumetrie vapour flow rate

V w * Vapour mass flow rate per unit area W Mass flow rate

W. .

Shaft work

U Overall heat transfer eoeffieient

U 0 . Overall heat transfer eoefficient based on tube outside area cjlv Volumetrie flow rate

Greek symbol

A

Latent heat

JlL Liquid viseosity

PL

Liquid density

Pv

Vapour density

cr

Surfaee ten sion

r

Tube loading Dimensionless numbers Nu Nusselt number Re Reynolds number Pr Prandtl number - 28-K K K K m3s -1 kg m-2 _s-1 kg s-1 W Wm-2_K-I Wm-2 K-I m3 _s-1 Jkg-1 Pas kgm-3 kgm-3 Nim kg m-1 s-1

(31)

Literature

1. Ikarashi, T., Chem. Econ. Eng. Rev. 12 (8), 31-34 (1980)

2. Morikawa, Y., Takagi, K., Moro-oka, Y., Ikawa, T., Chem. Lett. 1982,

1805-1808

3. Sodesawa, T., React. Kinet. Catal. Lett 24 (3-4),259-266 (1984)

4. Tonner, S.P., Trimm, D.L., Wainwright, M.S., Ind. Eng. Chem. Prod. Res.

Dev. 1984 (23), 384-388

5. Ai, M., Appl. Catal. 11,259-270 (1984)

6. Takagi, K., Morikawa, Y., Ikawa, T., Chem. Lett. 1985, 527-530

7. Sodesawa, T., Nagacho, M., Onodera, A., Nozaki, F., J. Catal. 102,460-463

(1986)

8. Guerrero-Ruiz, A., Rodriguez-Ramos, I., Fierro, J.L.G., Appl. Cata!. 72,

119-137 (1991)

9. Morikawa, Y., Adv. Catal. 39,303 - 327 (1993)

10. Ushikubo, T., Hattori, H., Tanabe, K., Chem. Lett. 1984,649-652

11. Taylor, D., Walker, D.R., J. Chem. Soc. (A), 1969,2991-2995

12. Lee, S.J., Lee, K.H, Lee, J.S., Appl. CAta!. A 83, 165-178 (1992)

13. Ma, F-Q, Lu, D-S, Guo, Z-Y, J. Molec. Cata! 78, 309-325 (1993)

14. Ma, F-Q, Lu, D-S, Guo, Z-Y, 1. Catal. 134,644-653 (1992)

15. Kim, K.M, Woo, H.C., Cheong, M., Kim, J.C., Lee, K.H., Lee, J.S., Kim,

Y.G., Appl. CAtal. 83, 15-30 (1992)

16. Sinnott, R.K., Coulson and Richardson's Chemical Engineering, Volume 6

(Design), Second Edition, Pergamon Press, New York, 1993, Chapter 12.

17. Stephanopoulos, G., Chemical Process Control, Prentice-Hall, Englewood

Cliffs NJ, 1984, Chapter 25.

(Design), Second Edition, Pergamon Press, New York, 1993, Chapter 6.

19. Monûoort, A.G., De Chemische Fabriek, Deel IT, Faculteit der Scheikundige

Technologie en der Materiaalkunde, vakgroep Chemische Procestechnologie, 1991, p. IIT-20

(Design), Second Edition, Pergamon Press, New York, 1993, p. 189.

Technologie en der Materiaalkunde, vakgroep Chemische Procestechnologie, 1991, p. ill-49

23. Process Economics Project, PEP Cost index update No. 46, SR! international,

Menlo Park CA, 1994.

24. Montfoort, A.G., De Chemische Fabriek, Deel IT, Faculteit der Scheikundige

Technologie en der Materiaalkunde, vakgroep Chemische Procestechnologie, 1991,p. ill-2

Technologie en der Materiaalkunde, vakgroep Chemische Procestechnologie,

1991, p. IIT-49 .

Technologie en der Materiaalkunde, vakgroep Chemische Procestechnologie, 1991, p. II-38

28. Process Economics Project, 1992 yearbook, Volume 2 (Germany), SR!

(32)

29. Dijkstra, L., Winkel, M.L., Kleinschalige Productie van Waterstof door Steamrefonning van Methanol, FVO 2949, Faculteit der Scheikundige

Technologie en der Materiaalkunde, vakgroep Chemische Procestechnologie, 1992,p.23

30. Process Economics Project, 1992 yearbook, Volume 2 (Germany), SRI

international, Menlo Park CA, 1992, p. 2.118.

31. Grievink, J., Meijer, F.A., Ham, A. van den, Handleiding voor het maken van

een Fabrieksvoorontwerp, Faculteit der Scheikundige Technologie en der

Materiaalkunde, vakgroep Chemische Procestechnologie, 1993, p. II-2

32. Raghuraman, K.S., Geels, H.P., Tio, T.H., Ratan, S., Kinetics Technology

Internationa11987 congress proceedings, Carbon monoxide production technologies, KT! BV., Zoetermeer, the Netherlands, 1987.

(Design), First Edition, Pergamon Press, New York, 1991, p. 459.

(Design), First Edition, Pergamon Press, New York, 1991, p. 488.

(33)

30-Hl H2 H3 H4 H5 H6 H7 H8 H9

!

-

---

--

-0-

-

--

-

---

-

---

-: "'- ... _Ycarbon methanol methanol HP HP steam purge T3 ( ) ( ) ( ) ( ) ( )

Hydrogen seperation Pyrolysis ofmethylformate Carbon monoxide seperation Dehydrogenation ofmethanol to methylformate methanollformate

distillation

Heat exchanger Rl Pyrolysis reactor _{PROCESS SHEET for the PRODUCTION} Heat exchanger R2 Dehydrogenation reactor

Vaporizer Tl Absorption column of high-purity CARBON MONOXIDE

Heater TI Absorption column

Cooler T3 Methanol / formate distillation column M. Kooistra Stream number FVO 3103

Heat exchanger Cl Compressor _{H.l. Meester} October 1994

Heater PI Pump _{B.F. Tuinstra} _{Temperature in K}

Heat exchanger P2 Recycle pump Reboiler VI Expansion valve

HlO Condenser Absolute pressure in bar

HP steam

>

"C "C ~ c:l,.

....