A new concept of a reactor sysstem for carrying out fast and highly exothermic reactions, in respect to the partial oxidation of o-xylene

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adres:

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F.V.O. Nr.

3040

Vakgroep Chemische Procestechnologie

Verslag behorende

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Faculteit der Scheikundige Technologie en der Materiaalkunde

Technische Universiteit Delft

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A new concept of a reactor

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stem

for carrying

out fast and highly

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exothemic reactions

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in respect to the partial oxidrtion of o-xylene .

Supervisor:

Dr. Ir.

A.

C

y

bulski

Students: J. Immerzeel Broerhuisstraat 40 2611

GD Delft

015-120853 January 11, 199~

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Gonfalone Maziestraat 20

2514 GT Den Haag

070-3606987

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Fabrieks voorontwerp: A new concept of a reactor system for carrying out fast and highly exothermic reactions.

Students: O. Gonfalone and J. Inunerzeel Supervisor: Dr. Ir. A. Cybulski

Table of Contents

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1. Sununary ... 2

2. Introduction ... 4

3. General infonnation ... 5

3.1. Production and products of phthalic anhydride ... 5

3.2. Developments in technology ... : ... 5

3.3. Present status oftechnology ... 6

3.3.1. Process conditions ... 6

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3.3.2. Reaction mechanism and kinetics ... 7

3.3.3. Process safety ... 7

3.3.4. Catalyst deactivation ... 8

3.3.5. Environmental aspects of the phthalic anhydride process ... : ... 9

4. Modeling ofthe two reactor system and cost estimation ... 10

4.1. Assumptions ... 10

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4.1.6. Assumptions conceming the modeis ... 10

4.1.7. Concerning both reactors: ... 10

4.1.8. Catalysts and catalyst deactivation ... 10

4.1.9. The multitubular reactor ... 11

4.1.10. The new adiabatic reactor ... 11

4.1.11. Costs assumptions ... 11

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_{4.2. The heterogeneous 2-dimensional tubular reactor model. ........... 12}

4.2.1. The concentration and temperature balances ... 12

4.2.2. Kinetic equations ... 13

4.2.3. Catalyst deactivation ... 14

4.2.4. Thennodynamic and physical properties ... 16

4.2.5. Numerical techniques ... 19

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_{4.3. Results of simulations .......................................... 19}

4.3.1. Operation conditions and model parameter values ... 19

4.3.2. Conversion and selectivity for conventional and the new reactor system ... 22

4.4. Costs analysis ... 23

4.4.1. Costs calculation ... 23

4.4.2. The safe distance between the adiabatic and the multitubular reactor ... 25

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_{4.4.3. Cost estimation of the new reactor concept... ..................... 26}

4.4.4. Profitability indices ... 26

5. Conclusions ... : ... 27

6. Recommendations ... 28

7. Nomenclature ... 29

8. References ... 31

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_{Al Chemical and thennodynamical properties ................... 33}

A2 Process safety, the Mond index and dangerous properties of the organic compounds ... 37

A3 Calculation of diameter of the pipes ... 43

A4 All results of simulation ... 44

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A5 Calculation ofthe heat losses between the multitubular and the adiabatic reactor... ... 49

A6 Calculation of costs of new concept ... 52

A 7 Pressure drop and reactor cost ... 60

A8 The PASCAL program ... : ... 68

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Fabrieks .... oorontwerp: A new concept of a reactor system for carrying out fast and highly exothennic reactions.

Students: O. Gonfalone and J. Immerzeel Supel"iisor: Dr. Ir. A. Cybulski

1. Summary

A new reactor concept is examined to increase the on-strearn time of a catalyst for the partial mddation of hydrocarbons. The new reactor concept consists oftwo reactors in series: (1) a multitubular reactor and (2) an adiabatic reactor. The system for carrying out the highly exothennic oxidation of o-xylene to phthalic anhydride is computer simulated. When catalyst activity in the first reactor is such that the conversion drops below acceptable minimum, a second adiabatic reactor is used to complete oxidation of o-xylene. The multitubular reactor was described by a two dimensional heterogeneous model. Orthogonal coUocation techniques were used to approximate the differential equations. The system of coup led equations was solved by Ne",ton-Raphson iteration. The adiabatic reactor was simulated using the same turbo PASCAL program. ~the radial dimension of the mod'èl

Was

-redundant.

The adiabatic reactor fiIIed ",ith spent catalyst is capable of multiplying the on-strearn time by a factor of almost four. During this e~1ension of the on-strearn time, the time averaged selectivity towards phthalic anhydride remains practically unchanged. During the extension the conversion of o-xylene is almost complete. The operating time ofthe new reactor concept is limited by the maximum permissible catalyst temperature.

Although the new reactor concept does increase the on-stream time, as intended, it is not economically

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This is caused by:

• The high installation co st of the new reactor concept. The cost are mainly determined by the costs of stainIess steel piping and appendages (88% of total cost).

• The cyclic operation. When the reactor is inactive, interest and depreciation form a large negative cash flow. When the reactor is operative then the ta"es consurne a large part of the profits.

The total investrnent ofthe new reactor concept is one miIIion Dutch guilders. The payout time is 7.7 years or 10.4

years when the electrical e~1ra energy consumption caused by the increased pressure drop over the reactor system is considered. The costs of catalyst replacement and the off-stream time are reduced by 59% and 37% without and including the ex1ra energy consumption respectively. The internal rate of the return is -9%.

The new reactor concept can ooly be economically viabie when the total installation costs are reduced.

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Fabrieks vooront\verp: A new concept of a reactor S)·stem for carrying out fast and highly exothennic reactions.

Students: O. Gonfalone and J. Immerzeel Supervi${)r: Dr. Ir. A. Cybulski

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Introduction

The objective ofthls fabrieksvooront\Verp is to evaluate the performance of a new reactor concept by modeling, by application ofthe concept for the production ofphthalic anhydride fiom o-x-ylene. Circumstances must be made clear to explain the contents and the practical value ofthis concept.

Partial oxidations of hydrocarbons are catalytic processes, which are carried out mostly in multitubular reactors.

Full conversion of hydrocarbons is usually desirabie for at least t\Vo reasons: (i) separation and purification of

products become more complex and ex-pensive if raw material is present in the mixture leaving the reactor, andlor

(ii) an unreacted raw material must be recovered or combusted. Both operations increase production costs. At the

initial stage ofthe reactor operation a temperature peak is relatively high and conversion ofthe raw material is practically 100%. A breakthrough of raw materials occurs after a certain period of the reactor operation due to

catalyst aging. Tben the plant is stopped, the catalyst is replaced and the plant is restarted. Tbe plant is out of

operation for several days. An extension of the period bet\Veen catalyst replacements would result in decreasing of

an arnount of the catalyst consumed per unit of product and improving the

ratio

of changeover time to tbe time on

strearn. Tbe importance of the ex1ension of on stream time can be understood on taking in account that the price of

the catalyst is usually 1-2% of the total value of the installation, while the costs associated ... vith the interruption of

the installation functioning and recharcll1ng ofthe reactors with a new catalyst portion are more than 2.5% (11.

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AIR 10 12 11 o-xylene

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Figure 1: The: flow diagram ofthe: phthalic anhydride: process, including the: adiabalic r=:tor. I1 is obvious thaI th.:re: should he: valv.:s in the: piping 10 and oul ofthe: s.:.;ond r.:actor, howcver th.:s.: are nol drawn for rcasons of c1arily and simplicity.

L.:gl!tld: (1) air blo .... ·er: (2) o-1\-yII!tle: pump; (3) air prehe:aler: (4) o-1\-ylcne: prehe:aler, (5) carburelor, (6) convl!tllional multitubular re:actor, (7) salt .:ooler: (8) gas cooler: (9) aft.:r coolo=r, (10) switch condo:nsors, (11) adiabatic pack"d he:d r=:tor, (12) raw PA to purification.

The above goal can be reached using a system of two reactors. as ShO\Yll in figure I. The first reactor is a

multitubular reactor with fresh catalyst, the second reactor is an adiabatic reactor fil!ed \~ith spent catalyst

\\ithdrawn after thepreceding period of operation. Until conversion of a raw mate rial is higher than the pennitted

limit onIy the multitubular reactor is operated. When the breaJ.. .. through takes place the adiabatic reactor is started

to complete the reaction or to reach the predetermined convers ion leve!.

An adiabatic reactor is an unconventional solution for oxidation reactions that are fast and highly exothermic. Tbe

adiabatic reactor would be operated here only for completing the oxidation process. The amount of heat generated

in this reactor \Vil! be smal!. Obviously, there will be a need to control the adiabatic reactor in case of significant process perturbations, which might result in a temperature runaway. To prevent or to e.,1inguish the runaway the reaction mix1Ure can be quenched with steam.

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Fabrieks voorontwerp: A new concept of a reactor sy·stem for carrying out fast and highly exothermic reactions. Students: O. Gonfalone and 1. Inunerzeel Supervisor: Dr. Ir. A. Cybulski

3. General information.

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1. Production and products of phthalic anhydride

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Table 1 shows that the world year production ofphthalic anhydride is still increasing. In Western Europe and Japan, over 60% ofthe phthalic anhydride is used in the production ofphthalic esters; in the United States this figure is around 55%. Othet applications for phthalic anhydride are the production of alkyd resins and unsaturated polyester resins. each representing around 20%, completed by the production of dyes and pigments.

T a bi e 1 E stimate year pro uctIon . d d 0 f pi t a h h

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IC an v n e h d 'd

Estirnated year production of production in phthalic anhydride 103 tons

1976 1509 (2) 1985 2100 (3)

1988 >2900 [41

1992 >3300

. The main application for phthalic anhydride is the production of phthalic esters for use as plasticizers. Plasticizers are aLLx.iliary agents used in the industria[ processing of plastics. By lowering the intennolecular forces between the molecular chains, they endow high polymeric substances with certain desirabIe physical properties, e.g., reduced britttleness, lower hardness and where necessary increased adhesion.

The phthalic esters are the most important industrial plasticizers and, in practice, almost all commercially

available aliphatic Or cyclic a1cohols are used in the production of phthalates. Practically 85% of the production of phthalic esters is employed as plasticizers for PVC. The remaining 15% is used in aLLx.iliaries for paints,

dispersions. cellulose, polystyrene and other polymers. Besides phthalic esters, the most important applications for PA are in unsaturated polyester resins. together with alkyd resins produced by reaction with polyhydric alcohols. These polymers are used principally as raw materials in paint manufacture.

Two decades ago the largest use ofphthalic anhydride

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as a constituent ofthe ethylene glycol-terephthalic acid polymer, sold under trade names of Dacron and Mylar [51. Other uses for PA. but of less importanee in tenns of quantity, are the production ofpigments, dyes and phthalimide. Phthalimide is used as a raw mate rial in the production of anthranilic acid, pesticides and phannaceutical products.

3.2. Developments in technology

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Phthalic acid was discovered by the French chemist Auguste Laurent in 1836. Commercial production of PA was taken up by BASF in 1872, by the oxidation of naphthalene with magnesium dioxide and hydrochloric acid to obtain the required base material for the manufacture of dyestuffs and later for phenolphthalein; however the yield was onIy 5 to 7%.

Phthalic anhydride is produced from both naphthalene and o-xylene. The use of naphthalene is especially high in Japan, atjust 40%, whereas in the USA and in some West European countries it is relatively low, since in these countries naphthalene is most commonly used in the manufacture of dyes and agricultural chemica.ls. Currently 80% of the world production of PA is based on o-:\ylene [61, which is due to the lower price of o-xylene, its easier transportation, higher selectivity towards oxidation and higher purity of the product obtained.

Oxidation of aromatics, including o-:\ylene and naphthalene, to carbo:\ylic acids can be carried out in both the liquid and the liquid- and gas-phase. Liquid-phase oxidation is generally distinguished by high selectivity at high conversion rates. The disadvantage is that down-strearn processing of the reaction products requires separation of the reaction solvent. In gas-phase oxidation, the co st of purifying the reaction products is lower. since there is no solvent recovery involved; a drawback is the reduced selecti"ity, because of the necessarily higher temperatures.

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Fabrieks voorontwerp: A new concept of a reactor ~·stem for ca.n;.ing out fast and highly exothermic reactions.

Students: 0. Gonfalone and J. Immerzeel Supervisor: Dr. Ir. A. Cybulski

OnIy gas-phase oxidation has proved commercially .... iable for the oxidation of o-:\.-ylene and naphthalene. A high )'ield "\\ith high gas throughput, controlled removal of the heat of reaction \,,;th optimal use of the process steam, the highest possible precipitation and recovery ofthe crude product and efficient purification ofthe product are

essential features of the industrial proces~es.

The fust catalysts used for o-xylene oxidation to phthalic anhydride were based on vanadium pentoxide melts [7].

Irrespective oftheir long-term e:\.l'loitation (more than 10 years) and the relatively high selecti .... ity ofabout 70% of the theoretical, these catalysts have essential disadvantage: they are not efficient at high volume rates of raw material. In 1968 BASF applied supported spherical particIe catalysts with a coverage based on V205-Ti02' The excellent properties of the catalysts allowed the introduction of o-:\.l'lene as a raw mate rial in the production of

phthalic anhydride. Work by several authors has shown that activity and selectivity are achieved when the vanadia

is present in the form of highly dispersed, amorphous species, rather than as crystallites. The Ti02 content in the conunercial catalysts varies between 60% and 98%. while the V202 content ranges from 2 to 15%. Compounds of antimony, rubidium, cesium, niobium and phosphorus are added to improve the selectivity.

The vanadia-titania catalysts are characterized by a relatively low specific surface area (1-2 m2/g). Industrial

catalysts are prepared by deposition of a thin catalyst layer on an inert, mostly nonporous support. The application of thin catalytic film is connected with the strongly exothermic character of the oxidation reactions. The passive

nonporous support absorbs the heat evolved and essentially improves the thermal conditions ofthe catalysts (8J.

Irregu]ar spheres as weil as ring shaped catalyst particles are used in industrial practice. The ring type particles do have the advantage of a lower pressure drop, the spherical particles do have better heat conducting properties. The

current service life of the catalyst is of the order of 2 to 4 years (9).

3.3. Present status oftechnology.

3.3.1.

Process conditions.

The conunercial oxidation of o-:\."ylene in fixed bed is nowadays carried out exclusively in multitubular reactors with a capacity of up to 50000 tpa; large scale reactors are fitted with 25000 tubes, each with a diameter of 25 mmo

Concentrations of o-:\.l'lene in the feed vary, concentrations up to 90 g m-3 (STP) are used in industrial practice

[IOJ

It is possible to operate existing older plants, which \Vere originally designed for operating loads of 40 glm3 (STP)

with o-:\"ylene loadings of 60-70 glm3 (STP) by employing modern catalysts. High loading is particularly desirabie

to improve the energy efficiency. An undesirable effect of higher o-:\.l'lene concentrations is that the amount of heat that has to be transported from the reactor tubes to the salt bath increases more than proportionally in relation to

the increase in o-.\l'lene concentration. As heat transfer coefficients do not increase appreciably this resuIts in

higher catalyst hot-spot temperatures. These higher hot-spot temperatures can not be compensated for by using a correspondingly lower salt bath temperature, since the temperature in the exit part of the reactor would become too

low for intermediate produets to oxidize further. Thus operation at high o-xylene concentration resu]ts in an

increased therrnal strain on the catalyst. Consequently the catalyst must be adapted to these conditions. The

attainable increase in loading depends primarilv on the feasibility of remO'v'jng safely the heat of reaction from the reactor. Very high loadings of up to 90 g m-3 (STP) can be made possible by recycling the exhaust gases, thus by reducing the o:\';'gen content.

The gas flow rates used in industrial reactors vary between 2.5 and 8 m3 h- I tube-I (STP). The reactor tubes are

2.7-3.5 m long. The in1et temperature ofthe o-xylene/air mi:\1ure depends on the efficiency ofthe heat exchanger

used. The minimal temperature is deternlined by the boiling point of o-:\."ylene (l~~0C).

The heat of reaction is removed by means of a eutectic salt bath of potassium nitrate and sodium nitrite (59%/41 %)

\\ith a melting point of around I~ IOC and is used to generate high pressure steam, up to 70 bar.

The temperature of the salt bath is usually in the range of 370°C to 390°C and limited to ~OO°C.

According to literature data it is now possible to achieve yields of 120 \\-1%. but it is generally accepted that reactor

yields exceeding 112% are impossible on industrial scale. In o-:\y'lene based processes, after allO\\-ing losses during

purification, )ields of 108-110\\1% may be expected. This is equivalent to 77 -79% of stochiometric yield.

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Fabrieks voorontwerp: A ne~' concept of a reactor system for ca.rT)ing out fast and highly exothermic reactions.

Students: O. Gonfalone and J. Inunerzeel Supenisor: Dr. Ir. A. Cybulski

3.3.2.

Reaction mechanism and kinetics.

The oxidation of o-xylene to phthalic anhydride is a complex reaction, which includes the participation of 12 electrons, detachment of 6 hydrogen atoms and incorporation of 3 oxygen atoms in the o-xylene molecule.

The data available on the mechanism of catalytic oxidation of o-xylene are insufficient. This is probably due to the complex composition ofthe reaction gases, which makes their analysis difficult. The most used kinetic model is proposed by Calderbank et al [11). lts advantages consist in the prediction ofthe correct temperature and concentration profiles corresponding to a large-scale reactor (12_{1. .}

1 " " " - - - 1

--t~---l

c§(

o

CH3

--+

CH3

l

/c'

/0

C 11

o

11

--+

rQYC0'o

~CO;'

Figure 2: A gen~aliz.:d r.:action sch=~ ofth.: oxidation of ()o:,:yl.:n.: into phthalic anhydrid.:. Diif~.:nt inv.:stigators do all sugg.:st diif.:rent reaction sch.:m.:'s. Th~ reaction shown in this pictur~ contains all r~actions st~s of 7 mod~ls. A revi~w on all kin.:tic mod~ls cart b.e found in phthalic

anhydrid~ from ()oxyl.:n~ 111.

Other products of the oxidation of o-:'I.'Ylene that are mentioned in the literature are: maleic anhydride, mono- and dimethyl maleic anhydride, benzoic acid, toluic acid, benzene, anthraquinone and tany polymerization products.

Of these only maleic anhydride is considered to be important. The formation of maleic anhydride is observed at very high conversion degrees only and is most probably aresuIt ofthe oxidation ofphthalic anhydride oxidation

[13J. Maleic anhydride decomposes when heated above 150°C [14J. Therefore by the application of suitable processing conditions, maleic anhydride can be obtained as a by-product in levels up to 4% [15J.

3.3

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Process safety.

The flash point of a mixture of air and 0-.'I.'Y1ene is .+56

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The explosive range for o-.'I.'Ylene in air extends from 44 g m-3 to 335 g m-3 (STP). When the limits of flammability are exceeded, it is essential to estimate the risks to the safety of the personnel and of the plants. It is impossible to eliminate the sourees of ignition of an explosive mixture but the equipment can be designed to be inherently safe, i.e. to \'.ithstand the highest possible pressure.

The probability of ignition is real, the auto ignition temperature of o-:'I.)'lene is '+30°C [16_1.

According to Handbook of chemicals production processes (17) self-ignition does not have to resuIt in a flare up.

The author mentions tllat the only result so far noticed is a relatively pronounced sharp drop in product quantity and quality. When the catalyst is recovered from the tubes, observed damage ranges from a merely blackening of the coating of the particles to a rnelting of the carrier bedies caused by localized heating. Sometimes the darnage is confined to a few centimeters, sometimes it reaches al most the reactor exit. This picture is consistent with self-igrûtion ranging from a so-called "cold flare" to a reallocal e:-.'plosion. An essential point is that in all cases of igrûtion so far noticed, it was quenched so rapidly by the heat capacity of the catalyst that no harrn

was

done to the tubes or other equipment.

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Fabrieks voorontwerp: A new concept of a reactor system for carT)ing out fast and highly exothermic reactions.

Students: 0. Gonfalone and 1. Immerzeel Supervisor: Dr. Ir. A. Cybulski

Because of the relatively low temperature of the first decimeters of the tubes this region acts like a flame arrestor. Additionally the equipment may be safe garded by pressure relief devices, 50 that pressure surges are relieved safely.

Care must be exercised in the design and operation of these reactors because fires can result fiom accidental contact ofthe salt with the organic materials [18]. The production of PA has caused many industrial explosions (19].

Not only o-x]'lene is explosive in the form ofvapor when exposed to heat. The by-product maleic anhydride is also ex-plosive. Both anhydrides are corrosive. Maleic anhydride has an auto ignition temperature of 4

noc.

The reported ex-plosion hazard data of PA in air vary significantly. Explosions can occur at concentrations below 100 glm3, depending on the impurities present. Recent incidents in production plants indicate that PA

concentrations exceeding 35 glm3 in the reaction product gas are capable of ignition, if heat transfer salts enter the reactor due to broken tubes [20]. In Sax's dangerous properties ofindustrial materials [21] it is only mentioned that PA forms a moderate e.'\-plosion hazard in the form of dust when exposed to flame. The lower and upper limits of explosivity ofphthalic anhydride are 1.7% and 10.4%.

3.3.4. Catalyst deactivation.

According to Ullman's encyclopedia o/chemica! techno!ogy (1992) (221, no commercial catalyst has been

investigated in depth and there is no literature data relating to aging phenomena. In an ex1ensive review on phthalic anhydride from o-x)'lene by Nikolov, Klissurski et al [IJ the causes and the importance of deactivation are given:

Investigations of the deactivation processes occurring during the catalyst ex-ploitation are of economic and scientific importance. These investigations permit choosing the exploitation conditions in such way as to ensure a longer catalyst life time. The high cost of tllese studies is tlle reason for performing them in large-scale laboratories only, owing to which there is not sufficient information in this respect.

One of the most important characteristics of the industrial catalysts for the preparation of phthalic anhydride is the maximum admissible working temperature. The size and temperature maximum as weil as the temperature profile are functions of the basic process parameters (temperature of the coolant agent, gas flow rate, initial concentration of o-x]'lene), on the one hand, and the irreversible change in the catalyst , on the other. In the patent literature it is stated that the maximum admissible temperature of the catalyst should not exceed 500°C.

The transformation of anatase into rutile is considered as the main reason for the irreversible deactivation of the V205-TiOZ catalyst. The anatase-rutile transition has a very low activation of IS kcal/mol, while in the absence of V205 this transition proceeds at a temperature above 850°C \\'ith an activation energy of 150 kcal/mol.

It is assumed that the addition of water vapor to the reaction mixture leads to an enhanced selecti'vity. On the other hand Nikolov and Anastasov [231 conclude that prolonged addition of water vapor to the initia! o-x]'lene/air mix1ure leads to an irreversible deactivation of the catalyst. This is ascribed to a dec rea se in the phosphorous content of the catalyst, as a result which its acid-basic characteristics are strongly deteriorated. A similar decrease in phosphorous content is obser'/ed after three years utilization under industrial conditions.

760

T

740 • 1 mon1h 720

T

7001

)~~

A 6mon1hs

6801

6601

• 15 mon1hs

640

1 0 50 100 150 200 250

Figu re J: Changes ofth~ t"'llp",atur~ frofil~ in th~ ractor tub~ during ~:. .. ploitation. Inkt o-:\yI.:n~ concentration is 40 g m·3; gas rat~ 4.5 m hol (STP). Th~ I~gend contains th~ ~:\l'loitation tim~ ofth~ catalyst.

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Fabrieks voorontwerp: A new concept of a reactor system for carrying out iast and highly exothermic reactions. Students: 0. Gonfalone and J. Immerzeel Supervisor: Dr. Ir. A. Cybulski

The onIy quantitative data on catalyst deactivation found in literature are given in figure 3 and in table 2. The data were obtained by Nikolov & Anastasov [7]. Measurements were do ne in a single stainless steel tube. The internal diameter and height of the tube were 25

mm

and 3500 mm, respectively. The industrial vanadia titania catalyst used had a composition of 93A \\'t % Ti02' 6.0 \\1% V205, 0.2% \\<1% P205, 0.1 \\1% W03 and 0.1 wt% Al203' The support consisted of porcelain spheres 6 mm in diameter. The reactor contained 1.13 I of the catalyst. The ex-periments were perforrned under a pressure close to atmospheric with a relatively high excess of air. After fifteen months of ex-ploitation, the hot spot chariges its position by 40 cm and its temperature by about SOK. X-ray phase analysis showed that rutile-anatase transformation occurred to a small ex1ent onIy in the hot spot zone at tem~ratures higher than 730K. The specific surface area of this catalyst decreases significantly fiom 9.0 to 2.6 glm ',whereas in the zone of lower temperature this value

was

8.0 glm2. No changes in the catalyst composition after its deactivation where observed. Hence the deactivation is mainIy characterized by a surface area loss, reading to a decrease of activity per unit mass.

The inIet and coolant temperature were not specified. The coolant temperature must have been close to 650 1(,

because of the height of the temperature profile after 1.50 meters that is reached after 1 month of exploitation. The inlet temperature has very little influence. The gas is heated up in the initial 5-10 cm of the reactor.

T a hl e 2 : Ch anges 0 f I se ectlV1tv an d conversIOn d unng ex-ploltatlOn . I . 0 f h t e cata I vs t.

ex-ploitation time selectivitv towards conversion of (months) phthalic anhydride (%) phthalide (%) o-xylene (%)

1 73.6 0.005 99.9

6 72.8 0.10 99.8

15 69.6 0.67 99.7

Doubtless, a precise simulation of the oxidation process needs the consideration

of

the non uniform catalyst activity along the bed. The catalyst will be deactivated mainIy in the hot spot region. The activity of the catalyst depends on its catalytic surface. Unfortunately there are no suitable relationships between bath parameters, which can be considered in the ma thema ti cal model of the process. Therefore the catalyst activity is corrected by a constant factor. The value of the correction factor of fresh catalyst is 1. The value of the correction factor. as a function of the axial place coordinate, can be determined only by minimizing deviations between experimental and the results predicted by modeling [241. Because no experimental data is available, for this work the deactivation correction factor is assumed to be constant along the whole bed length.

3.3.5. Environmental aspects ofthe phthalic anhydride process.

When the activity of the catalyst becomes to low, it must be replaced. Used catalyst may be deposited, for an example, at a secure landfill site. This may depend on local regulations o\\ing to the toxicity ofthe additives. As a result of various problems. \'lashing and reusing of the active material has so far been uneconomical. even though proposals are repeatedly being made in this connection [25]. The intention of the new reactor concept. reuse of spent catalyst and ex1ension of service life of the catalyst. is undoubtedly an improvement out of environmental viewpoint.

The main substance emitted from phthalic anhydride plants is carbon monoxide, which cannot be removed by exhaust gas scrubbers. A plant producing 10000 tJa and operating \\ith a loading of 60 g m-3 (STP) o-xylene in the process air emits ca 1500 tJa carbon monoxide. depending on the operating conditions. Owing to gro\\ing

emironmental sensitivity is expected that in future it \'viII onIy be possible to omit the exhaust gasses after treatment in a purification plant. The exhaust gases can be c1eaned either by thermal combustion at 650-850°C or by catal)tic combustion at 250~50°C. The new reactor concept \'álI reduce the amount of unreacted hydrocarbons but \'viII not reduce the emission of carbon monoxide.

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Fabrieks voorontwerp: A new concept of a reactor SJ'stem for carrying out fast and highly exothennic reactions.

Students: O. GonfaJone and J. Immerzeel Supervisor: Dr. Ir. A. Cybulski

4

.

Modelïng of the two reactor system and cost estimation.

DL

6

€-k.a

v

i

(.?

\{..~

4.1. Assumptions.

7T

(.? 'N _ l

~

'" c.t-\.N"')"Ç 0 'r

t

'P

~

el-t

0

~t..eLe

/

*'

r-~I.(.(Á

_

UL-vh>r

1..

4.1.6. Assumptions concerning the modeis.

*

Cl..

~

\

J

~

pt-f')

1

~cl

...

bI

Ul:>~

bL

CO"'-) I

era-

e

• Herce-Virgil et al. [26] have simulated the oxidation of o->.:ylene to phthaJic anhydride using one- and

two-rumensionaJ modeis. The verification of the caJculated temperature and concentration profiles is achieved by

experiments in a pilot reactor. It is found that the best description of the ex-perimentaJ profiles is made by the

climensional model. Therefore oxidation of o-xylene in the multitubular is simulated for a

two-dimensional heterogeneous model. .

• The adiabatic reactor can be simulated by the same model when radial transport coefficients are made equal to

zero. The model is time dependent to simulate responses of possible temperature runaways on steam quenches.

4.1.7. Conceming both reactors:

• It is assumed that the ideal gas law may be applied to estimate some of the properties of the gas phase. This

assumption is justified by the low pressure.

• The properties of the gas in the reactor are taken as that of air. The influence of the organic compounds and

combustion products are neglected. The simplification is justified by the value of the inlet concentration of

0-xylene of -l0 gm-3 (STP), equal to 31 g o-x-ylene per kg air.

• The diffusion coefficients of the components in the reactor differ in such a way that an average value can

represent all diffusion coefficients.

• After the reactor(s) the pressure of the gas containing raw PA is 1.2 bar. This pressure should be enough for

the gas to flow through the purification section of the plant.

• The density of the gas is proportional to the pressure and inversely proportional to the temperature of the gas.

The density ofthe gas and thus also the superficial velocity depend on temperature. These variables influence almost every parameter in the mathematical model. Therefore all parameters are calculated for temperatures in the range of375°C to 475°C, and their temperature dependency is examined.

• The pressure dro . s neolected. The pressure in the reactor is the average of the in- and outlet

pressur . Therefore the pressure drop does not resu 111 an lIlcrease in tempera tu re.

?

---4.1.8. Catalysts and catalyst deactivation.

• The catalyst is spherical wilh a diameter 3 millimeters. Beside this particIe size also particles of 6 mm are used

in ex-periments described in literature.

• Because of the lack of detailed deactivation data, the deactivation can not be simulated as a function of time,

place and temperature. The deactivation of the catalyst is ass d second order \"ith the activity of the

catalyst, and constant along the whole bed length econd order deca)" is minimu higher order of decay

\\ill resuIt in a steeper initial decay but a higher acti"ity in when the time increases.

7

• The catalyst can be used for two years before it must be replaced.

• The con ... ersion of o-x-ylene at the end of the catalyst life is 99.5%, when the available deactivation data is

considered this value seems reasonable.

• The catalyst deactivates unifomlly along the whole bed length .

• The activity of the catalyst in the adiabatic reactor decreases as fast as the activity of the catalyst in the

multitubular reactor. This assumption must be made because no deactivation at all is less realistic. Because of the usually lower maximum temperature in the adiabatic reactor, the deacti\"ation in this reactor is

exaggerated. The exaggeration has little influence the conversion, the temperature rises until almost all is

converted .

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Fabrieks voorontwerp: A new concept of a reactor system for carrying out fast and highly exothennic reactions.

Students: O. Gonfalone and 1. Immerzeel Supervisor: Dr. Ir. A. Cybulski 4.1.9. The multitubular reactor.

• The hot-spot temperature is limited to 460-470 oe. It is advised to limit the temperature to 500°C, but what must be considered is that the cross sectional temperature profile decreases from its maximum in the middle to a few degrees above the coolant temperature at the tube wal!. The catalyst temperature in the middle ofthe reactor tube is always above the average temperature.

• The wall temperature is initially equal to 370°C, to maintain constant selectivity and conversion the

temperature is gradually increased to 390°e. The same wall temperature of390°C is used during operation of the new reactor concept.

4.1.10. The new adiabatic reactor.

There are many degrees of freedom in the design of the second reactor. The foJlowing assumptions are made: • The ex"tra pressure drop over the new reactor and newly installed pi ping is such that it can be covered by the

surplus capacity al ready existing air pump.

• The amount of catalyst in the adiabatic reactor is equal to the amount of catalyst in the multitubular reactor.

When the catalyst in the first reactor is deactivated in such degree that it becomes economically favorably to replace it, then the whole amount of catalyst is placed in the new adiabatic reactor. The whole amount of catalyst is reused to ma.ximize the degree of conversion and to increase the maximum operating time ofthe adiabatic reactor.

• The catalyst in the adiabatic reactor deactivates by the same mechanism and as fast as the catalyst in the multitubular reactor.

• The minimum height of the catalyst bed in the adiabatic reactor is 1.5 meters. A lower bed height would imply a lower pressure drop, but can also cause maldistribution of the fluid.

• The temperature loss between the outlet of the multitubular reactor and the inlet of the adiabatic reactor is reduced by pipe insulation. A loss of 8 Kel .. in is assumed between .the outlet temperature of the multitubular reactor and the inlet temperature of the adiabatic reactor. To much temperature loss would implicate no or almost no reaction in the adiabatic reactor. Calculations in appendix 5 show that this assumption is a safe exaggeration. Reheating the gas by a heat exchanger wiJl increase operating costs and cost of investment in such an ex1ent that the new concept becomes very unattractive.

• The use of the adiabatic reactor is limited to the ma."imum allowed temperature of the catalyst (500°C). When the pipe between both reactors is perfectly insulated then thls implies that the allo\ved temperature rise over the new reactor is approximately 100°C. Figure 13 illustrates that this roughly 10% of unreacted o-xylene can be safely converted to phthalic anhydride. The other reactions do have a much lower heat of reaction.

• The starting up of the adiabatic reactor requires no extra equipment. The gasses entering the reactor wiII first heat the reactor up, when its temperature is high enough the reactor starts to operate.

4.1.11. Costs assumptions

• The Lang factor used for the capital investment is based on the direct costs and indirect -costs that are relevant for this particular process.

• The price of the insulation is based on the economical insulation thickness.

• The cost of the reactor is based on the price of stainless steel.

Because the real price of the catalyst is not available, the price of the catalyst is estimated. Two statements were found in literature containing infonnation on catalyst price.

• The cost of the catalyst is 1-2 % of the total cost of the installation and that the costs of recharging and the off-stream time are more than 2.5 % of the total cost [11. The contribution of the 10ss of profits because of the off-stream time is unknown.

• The cost of catalyst and chemicals is 10.0$ /ton product produced by a 60 000 ton/year plant build in 1983 [271.

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Fabrieks voorontwerp: A new concept of a reactor system for carT)ing out fast and highly exothermic reactions.

Students: O. Gonfalone and 1. Immerzeel Supenisor: Dr. Ir. A. Cybulski

4

.

2. The heterogeneous 2-dimensional tubular reactor model.

4.2.1.

The concentration and temperature balances.

The concentration and temperature balances on the flo\\-ing fluid in a unit volume of reactor are given by:

éJc

l l

-éJz

I,'

éJc

+&-éJt

1

éJ

éJc

- DER--(r-)

r éJr

éJr

:.r

l!J

d

u.~

"<)"=( ')Lc..

7

(3 -

IDl:~

'0

~"2

' I.:

-k

_g

a

_v

(c -c)

_S

=0

(eq 1) (eq 2)

The concentration balance is valid for each component. The balances are based on the balances in Chemical

Reactor Analysis and Design [28]. The balances are slightly modified. Originally not all tenns had consistent units and the model was time independent. Now the model can also be used to simulate the non-stationary behaviour of

the responses of second reactor on steam injections. It is aIso possible that multiple steady states exist in the first

b

.

C "

1><

reactornon-stationary model is ab Ie to ca\culated the , the one that is significant is detennined COIT

bY~

steady stat . .. onditions. In case of multiple steady states only a

The temperature and concentration balances on tlle cat y t aTe~given by:

d

dt

éJc

_s

ot

_:,r

:.r

The mathematical model is accompanied by the initial and boundary conditions:

atz=OthenC=Co ,

T=~.o

andTs=7;

.

o

atr

=

0 then

oC/éJr, éJTs

/

or and oT

/

or

=

0

atr=Rthen

a~CT",-T)=À~oT/or,

a:CT",-1"s)=ÀS.oTs/or and

éJC

/

or=O

~_ w /~ ~

The imaginary time variabie , is defined to simplify the mass and temperature balances:

(eq 3)

(eq 4)

r

=

t - Z

&/u

__

~ (eq 5)

The second tenn of equation([ijbecomes negible when the time since startup increases and the gas velocity is

sufficiently high. For an)' dependent variabie F equatiOrtYlO@ re valid:

\\01.,'(

éJF

éJF ot

oF

0::

oF

-

= - -

+ -

-

= -

(eq6)

éJr

éJt

or

êz éh

ot

: I

The balances on the fluid phase can be simplified into:

êlo:.

:1

oz:

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Fabrieks voorontwerp: A new concept of a reactor system for carry;ng out fast and highly exothennic reactions. Students: O. Gonfalone and J. lnunerzeel Supenisor: Dr. Ir. A. Cybulski

oe

u

-oz

1

0

oe

- DER--(r-)

r

or

or

-k

g

a (c -e)

v S

=

0

r.r r,z r.r r.:

The balanees on the catalyst phase become:

c

oTs _

À,s

.!.~(r oTs)

Ps Ps

or

ER r

Or

Or

Z.T r.: (eq 7) (eq 8) (eq 9)

oc

s

or

_=.r (eq 10)

The model can be used to describe both reactor types. In case of the adiabatic reactor all radial profiles can be

neglected, only the first boundary condition is still appropriate.

4.2.2. Kinetic equations.

The kinetic model as proposed by Calderbank et al. [11] is drawn in figure ~, whereas A, B, C and D are o-xylene, o-tolualdehyde, phthalide and phthalic anhydride. Erepresent both carbon monoxide and carbon dioxide. Step 6 is estimated as kinetically insignificant.

k4

I

k

₁

k

₂

k

1

A

B

C

5

·n

I

k~

•

E

k6

I

Figure 4: Th" kinetic mod,,1. wher"as A. B. C and 0 are G-xyl.:ne. G-loluald"hyd". phthalide and phthalic anhydrid". E f.:pf=l both carborunonoxid" and carbondioxid".

The reaction equations are:

CSHIO

+

02 CSHSO

+

02 CSHIO

+

9.5 °2 CSHIO + 3 °2 CS~02 + 02 +H20 +H20 5 H20 + 3 H20 +H20 +2 CO (r I) (r2) (r3) (r4) (r5)

For this work it was assumed thaI the carbon oxides are fonned in the ratio of 1:3. This ratio was also used in the work of Kershenbaum and Lopez-lsunanza [29]. Nik%v et al [301 detennined that trus ratio starts of \,';th I: 1.8 \\ith fresh catalyst and increases to 1:3.1 during e:\-ploitation of the catalyst.

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Fabrieks voorontwerp: A new concept of a reactor s:-'stem for ca.rrying out fast and highly exothennic reactions. Students: O. Gonfalone and 1. Immerzeel Supenisor: Dr. Ir. A. Cybulski

The kinetic equations in partial pressures are:

'i

=kJ

a~ r_z

_=kz

alt

r

3

=k3

aPa

r~

=k4

aPa

rs=ks

a~ with (eq 11) (eq 12) (eq 13) (eq 14) (eq 15) (eq 16)

The values of the parameters in equations 11 to 16 are given in table 3, these values are obtained from Calderbank et al (111. The value ofKJ>ox with air at near atmospheric pressure is 0.00772 mol kg-1çl.

T a e bi 3 N umenca va ues . 1 0 fth e parameters 0 fth A h ' e rr enllls reactIOn rate equatlOn.

1 2 3 4 5

Eai 61420 46473 51204 54512" 57945

[I/mol]

ko,i 0.03778 0.00549 0.00353 0.01278 0.03148

[mol/(kg sPa)]

The reaction rate constants ki are regular Arrhenius type rate equations. The partial pressures can be easily transformed into concentrations by application of the ideal gas law. The Arrhenius reaction rate equation is modified to obtain :

-Ea

k

=

R T

ko

exp(--')

I all .1

R T

all 4.2.3. Catalyst deactivation (eq 17)

At page 9 it is mitten that the loss of activity is mainly due to the loss of surface area. Sintering also referred to as

aging, is the loss of catalytic acti,;ity due to a loss of active surf ace area resulting from the prolonged exposure to high gas-phase temperatures. The catalyst deactivates by sintering. The basics on sintering can be found in

Elements of chemical reaction engineering [31

J

.

For deactivation by sintering, the rate of deactivation is independent on the mean stream concentration. A1though other forms of the sintering decay law exist. one of the most commonly used decay laws are second order with respect to the present activity.

,

dl]

r

_d

=

kd

TT

=

-dl

Integrating. with Tl= 1 at time t=O, yields

(eq 18)

(eq 19)

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Fabrieks voorontwerp: A new concept ofa reactor system for carrying out fast and highly exothermic reactions.

Students: O. Gonfalone and 1. Immerzeel Supervisor: Dr. Ir. A. Cybulski

The sintering decay variabIe kd follows the Arrhenius equation:

(eq 20) The sintering becomes a function of place (two-dimensional) and the history of this place in the reactor. The loss of activity is directly proportional to the 1055 of surface area. The decrease in surface area of silica-alumina

can

he described by a power law fit:

dS =-kS"

dl

(eq 21)

For any order of n> 1 the following equation can be derived when equation 21 is properly integrated with the

~ppropriate boundary conditions:

(eq 22)

The order of surface decayof spherical silica alumina particles is 2 to 13. In figure 5 the catalyst activity is plotted against time for thirteen different orders of decay. After two years the catalyst life time is increased by the order of decay. The second order decay imp lies shortest possible catalyst life, the assumption of a second order decay im lies a conservative catalvst life-time.

100 90 80 70 ~

.,

60 ~ ₅₀ ;..

-

:.J 40

:

~

=

1~

..

""

"

",

.,

~"",,~,

... 1 0.0 1.0 2.0 3.0 4.0 5.0 6.0 7.0 8.0

time of catalyst utilisation [year]

Figure 5: Catalyst activity Tl as a function of firn 10 thiru=th ordo:T of d"cay. Th" activity al th" start is "qua I 10 100% and afkr 2 y.:ars tbe:

activity is assum"d 10 bc: "quailo 99.5%. for all ord= of d.:cay.

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Fabrieks voorontwerp: A new concept of a reactor Sj·stem for canying out fast and highly exothermie reactions.

Students: O. Gonfalone and J. Immerzeel Supervisor: Dr. Ir. A. Cybulski

4.2.4. Thennod)'namic and physical properties.

In the models ofYagi and Kum, Kuni and Srnith and later Sehlunder the effeetive thennal eonduetivity in the radial direetion is considered to consist oftwo contributions, the first dynarnic (dependent on flow conditions) and the seeond statie, so that:

(eq 23) The statie eontribution is build up from rt throu h the fluid in the voids and tran rt in whieh the solid phase is involved. e ehlundeiarrived at the following fonnula for the statie eontribution:

[

]'.

(eq 24)

where À-s and À-g are the conduetivity's of the solid and the gas phase respeetively, B,

e

and ars are deseribed by the folIO\ving equations. The radiation coefficient of the solid !.i/(m1 _{s K)] is given by:}

a

=0

.

227~(T+273)3

rs 2 -

P

100

(eq 25)

where p is the emrnissivity of the solid (=0.8) and T is the temperature [0C]. B is a funetion of porosity only:

B

=

1.25[(1-

&)/&]1019

(eq 26)

e

is the complex function :

(eq 27)

The dynanuc contribution arises exciusively from the transport in the fluid. The dynarnie eontribution is described

by: - - - -where: q' _ _

1 _ _ _ _

0_

.

1_4_

- Pemr - 1

+46(d

p

)2

de

16 -(eq 28) (eq 29)

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Fabrieks voorontwerp: A new concept of a reactor SJ·stem for canying out fast and highly exothennic reactions.

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The radial etfective heat conductivity of the gas and the solid phase are approximated by the dynamic and the static contribution respectively.

..1!ER

=).!

ER

l~ =l~

(eq 30 and 31)

The wall heat transfer coefficients are calculated according to De Wash and Froment [32_{1, by application of}

equations 32 and 33:

(eq 32 and 33)

A serie of dimensionless numbers are needed as parameters in equations for the estimation of the diffusion, heat and mass transport parameters.

Re=Pg ud/7]g

Re. =4uPg /(a

v

7]g)

Pr

=

7]g Cp/lg

Sc= 7]g/(D", Pg)

(eq 34, 35, 36 and 37)

The coefficients DER' hf, ha and Cl._{w in equations 13}, I .. , 15, 16 of the model are obtained from the following

equations, which are copied from a paper by Anastasov, E/enkov and Nik%v (331.

DER

=

Dm

(0

.

28

+0.053Re.

Sc)

a

_...

=

0 180

_. À. _g

ReO.

8

/d

_p

kg

=

0

.

357v,,/(Re0

36

Sc

0.67

E)

hl

=

0

.

384v" Cp/(Pg

ReO

.

36Sc067 E)

(eq 38, 39,40 and 41) In equation 38 Dm is mean coefficient of molecular diffusivity. This coefficient can be estimated by application of Fuller's method [34

1

.

The voidage of the catalyst packing in the reactor is given by:

(eq 42) The average outside surface area of the catalyst particles can be calculated \\ith equation .. 3:

(eq 43)

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Fabrieks voorontwerp: A new concept of a reactor system for carrying out fast and highly exotbermic reactions. Students: O. Gonfalone and J. Imrnerzeel Supenisor: Dr. Ir. A. Cybulski

Tbe heat of reaction of each reaction can be calculated witb:

è:J.f:

=

t

(v;

(Mi~~8

+ CpA; (T - 298)+

C~B;

(T -298)2 +

C~C;

(T - 298)3+ C:D; (T -298)4»

... (eq 44) In this equation ui is tbe reaction coordinate of component i in tbe reaction equations (R l-RS), tbe temperature T is in [0C].

The equation used most to calculate pressure drop in a packed bed is tbe Ergun equation. The effect oftbe temperature profile and tbe change of the number of molecules is neglected for simplicity. A modification of tbe Ergun equation for variabie density is given by (351:

p

=

Po

(1-

2/30

LIP

JO.s

(eq 45)

In equation 45 Po is tbe pressure ofthe gas [lbf'ft2), before entering the bed, P is the pressure ofthe gas [lbf'ft2) that passed the bed, L the length ofthe bed [ft) and ~o a factor [lbf'ft3) obtained from equation 21:

(eq 46)

In which G the superficial mass velocity [lb m/(ft2 s)), ~ a conversion factor 32.17~ [lbql ft/(s2 lb[)),

l1g the viscosity [Ibm I(ft _{s)). dp the partic\e diameter [ft) and} Po the gas density [lbn/fe) before entering the bed. The void fraction of the packed bed can be described by the general forrnula by Haugey and Beveridge (36J for spherical particles. The catalyst is assumed to perfectly spherical with a diameter of 3 mm. The void fraction plotted against the ratio of tube/partic\e diameter gives figure 6.

For relative large partic\es in a smal! tube the void fraction becomes 0.45. This void fraction \"il! be used in calculations of parameters and simulations of the multitubular reactor. On the other hand for relative small partic\es in beds with a large diameter the void fraction tends to 0.38. This void fraction will be used in calculations of parameters and simulations of the adiabatic reactor.

0.45

T

0.44 1 0.43

I

0.42

I

E 0.41

I

0.40

T

0.39

f/

0.38 0 0.1 0.2 0.3 0.4 0.5 d i a me te r p a rti c I e I d i a me te r tu b e

Figurc 6: The void frnction of a packed bed \\ith spherical partic\es as calculated by the forrnula by Haugey and Beveridge. plotted against the tube particJe diameter ratio.

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Fabrieks voorontwerp: A new concept of a reactor system for carry;ng out fast and highly exothermic reactions.

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The density, viscosity, conductiv;ty and heat capacity of air are copied from tab les in Transport Phenomena data companion [37]. The data are fit into simp Ie functions oftemperature. Thennod)llamic and physical properties of

o-:-.-ylene, carbon monoxide, carbon dioxide, water, nitrogen, o:-.-ygen and phthalide can be found in The properties

of gases and /iquids [S 1). The boiling points of o-tolualdehyde and phthalide can be found in Handbook of data on organic compounds volume 11 [481. These boiling poirits are used to estirnate the critical constants of o-tolualdehyde

and phthalide with Lydersen's method (381. The diffusion coefficients ofthe organic compounds are estimated

according to Fuller et al. (391. The standard enthalpies offonnation of o-tolualdehyde and phthalide are estirnated

with the group contribution method of Joback [5 1 I. The standard heat capacities of phthalide and o-tolualdehyde are

calculated with the method of Joback and with the method of Rihani & Doraiswamy [40

_1.

_{Results of all these}

estimations and calculations can be found in appendix 1.

4.2.5. Nurnerical techniques.

The model can be simulated on a computer when suitable solutions for the partial differential equations are known. An exact solution is impossible, the solution must be found nurnerically. In comparabIe studies on partial mddation of o-xylene in multitubular reactors the orthogonal collocation method is used to find approximate solutions for the

differential equations. The collocation method for differential equations is a trial-function technique. lts accuracy is

comparabie to the Galerkin's method. It goes far beyond the purpose ofthis paper to discuss all details ofthis

technique, therefore we refer to Si"1rl(Iation. regression and con trol of chemical reactors b y collocation techniques

[41

1.

-)

~ i~/.

C~~k

)\..{...{)

-~C)V\)

c..ll&t

,/el\.IL

ö~+e-t-~

/)()f1\..Vetre.

~

4

.

3. Results of simulations.

4.3.1. Operation conditions and model parameter values.

The temperature in the multitubular reactor can have a value in the range of 6J3-7~3K. The value of the

temperature dependent variables used in the simulation of the reactor is the average value of the variabIe over this

temperature range. A gas inlet temperature of 493 K is used. The inlet gas temperature is not a critical parameter

of the model because the gas is heated up to the wall temperature in the first centimeters of the multitubular reactor.

There are multiple operation conditions for the multitubular reactor. The multitubular reactor is simulated with

and without the use of the adiabatic reactor. The operating pressure over the multitubular reactor must be increased

to compensate the e:\Lr"a pressure loss due to the use of the adiabatic reactor. The pressure losses over the reactors

are calculated in detail in appendix 7. The operation conditions are summarized in tables 4-12. The te:\1S above the

tab les do contain all infonnation to c1arify their contents.

Table 4: Conditions of the multitubular reactor (reactor 1), the adiabatic reactor (reactor 2) does not operate. The

inlet pressure is 1.75 bar. the out let pressure is 1.20 bar. The pressure in the table is the average pressure.

L Length tube 3 [mI

[mi [m:l] [bar] [Nm:l/h] [Nm:l/s] [Nm/s] [mI [kg/m:l] [kg/m:l] [m2_/m:l] [m:l/m:l] 19

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Fabrieks voorontwerp: A new concept of a reactor S)·stem for carrying out fast and highly exothermic reactions.

Students: O. Gonfalone and 1. Inunerzeel SupelVisor: Dr. Ir. A. Cybulski

Table 5: The variables !hat are dependent on pressure, voidage and temperature. Average values are used in

. ul' fth I . fth I' bul ('th diabatic reactor).

sun auon 0 e re!rn ar operatIOn 0 emu tltu ar reactor W1 outa

T u _Pg Sc _DER kg (K) (m/s) (kg/m3_{] ( -) (m}2_{/s) [mis)} 625 2.125 0.832 1.40 4.171 E-4 0.0212 650 2.210 0.801 1.39 4.340E-4 0.0223 675 2.295 0.773 1.38 4.510E-4 0.0235 700 2.380 0.746 1.38 4.679E-4 0.0246 725 2.465 0.721 1.37 4.849E-4 0.0257 750 2.550 0.697 1.37 5.019E-4 0.0269 average 2.338 0.762 1.38 4.595E-4 0.0240

Table 6: This table contains the variables that depend on pressure, voidage and temperature. Average values of these variables are used in simulation of the operation of the multitubular reactor, while the adiabatic reactor is also operative. The inlet pressure of the muititubular reactor is 1. 91 bar, the outlet pressure is 1 A2 bar. The table bie ow ' IS ca cu ate or an averagepressure I I d ti 0 f 1 669 b ar. T u _Pg Sc _DER kg [K) [mis) [kg/m3_{) [ -) (m}2_{/s) [mis)} 625 1.879 0.941 1.2338 0.000370 0.02038 650 1.954 0.906 1.2282 0.000385 0.02145 675 2.029 0.874 1.2232 0.000400 0.02253 700 2.104 0.844 1.2190 0.000415 0.02362 725 2.180 0.815 1.2155 0.000430 0.02472 750 2.255 0.788 1.2128 0.000445 0.02582 average 2.067 0.861 1.2221 0.000407 0.02309

Table 7: In this table the variables that are independent on pressure or porosity are given. they do depend on temperature. Th e average va I ue 0 f h tese vana es are use . bi d' In t e slmu atlOn h . I ' 0 f the mulutubular reactor. .

T _"g À._g

C

_pg Re _{Reeq Om} Pr _hf a_w [K) (Pa*s) rW/fm Kl) [J/(kg Kl) [ -) . (-) (m2_{/s) [ -) [W/(m}2 _{u!(s m: 0C)]} °C)) 625 3.119E-5 0.04862 1061 170.14 206.24 2.685E-5 0.681 326.15 177.68 650 3.201 E-5 0.05008 1066 165.97 201.18 2.876E-5 0.682 330.80 179.42 675 3.284E-5 0.05151 1071 162.00 196.36 3.072E-5 0.683 335.26 181.00 700 3.366E-5 0.05292 1077 158.18 191.73 3.274E-5 0.685 339.50 182.44 725 3.449E-5 0.05430 1082 154.51 187.29 3.481E-5 0.687 343.53 183.71 750 3.532E-5 0.05566 1087 150.97 183.00 3.694E-5 0.690 347.34 184.83

average 3.325E-5 0.05218 1074 160.30 194.30 3.180E-5 0.684 337.10 181 .51 Table 8: In trus table the results of caIculations \\;th the Zehner and Schlünder (1972) theory are given. These results are an intennediate calculation result and are not used in the reactor simulation.

T À._s )-g Ct._rs B

e

\f' (K] [W/(m Kl] [W/(m Kl) (W/(K m2_{1] [ -) ( -] [ -]} 625 1.163E-3 4.862E-5 221.68 1.562 11216 0.08528 650 1.163E-3 5.008E-5 249.36 1.562 12248 0.08528 675 1.163E-3 5.151E-5 279.25 1.562 13335 0.08528 700 1.163E-3 5.292E-5 311.44 1.562 14477 0.08528 I 725 1.163E-3 5.430E-5 346.02 1.562 15675 0.08528 750 1.163E-3 5.566E-5 383.06 1.562 16931 0.08528 20

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Fabrieks voorontwerp: A new concept of a reactor Sj·stem for carrying out fast and highly exothermic reactions.

Students: O. Gonfalone and 1. Immerzeel Supervisor: Dr. Ir. A. Cybulski

Table 9: Radial heat transport coefficients in multitubular reactor, these variables are used in the simulation ofthe multitubular reactor. The coefficients do not depend on pressure. The adiabatic reactor does not require any radial transport parameters. T Îl.s_{ER ÎI.!ER Îl.} ER o.S w o.tw [K] [W/(m Kl] [W/(m KlJ [W/(m KlJ [W/(Km2_{l) [W/(Km}2_)] 625 0.0787 0.4802 0.5589 25.03 152.64 650 0.0884 0.4832 0.5716 27.75 151.66 675 0.0988 0.4861 0.5849 30.58 150.42 700 0.1101 0.4889 0.5990 33.52 148.91 725 0.1221 0.4916 0.6137 36.56 147.16 750 0.1350 0.4942 0.6292 39.67 145.16 average 0.1055 0.4874 0.5929 32.19 149.33

Table 10: This table contains the conditions ofthe adiabatic reactor. The pressure of 1.31 bar is the average pressure over the inlet pressure (1.42 bar) and the outletpressure (1.20 bar).

P pressure 1.3124 [bar]

cp gas flow gas flow A

u

total cross seetional area superticial gas velocity Dp diameter catalyst partieles Pcat density

Pbed density

specific external surface void traction 25000 [Nm3_/h) 6.944 [Nm3_/s] 9.079 [m2 ] 2.01 [Nm/s] 0.003

rml

2776 [kg/m3_] 1527 [kg/m3_] 1240 [m2_/m3_] 0.38 [m3_/m3 ]

Table 11: The variables used for the simulation ofthe adiabatic reactor, these variables are dependent on pressure, porosity and temperature. The adiabatic reactor is onlv used in combination with the multitubular rea ctor.

T u _Pg Se DER kg hf

[K] [mis] [kg/m3_{] [ -] [m}2_{/s] [mis] [W/(m}2 _°elJ

625 1.334 0.740 1.569 0.000236 0.017985 265.917 650 1.387 0.713 1.562 0.000245 0.01893 269.710 675 1.440 0.687 1.556 0.000255 0.019884 273.341 700 1.494 0.663 1.550 0.000265 0.020847 276.801 725 1.547 0.641 1.546 0.000274 0.021816 280.086 750 1.600 0.620 1.542 0.000284 0.022792 283.195 average 1.467 0.677 1.554 0.000260 0.020376 274.842

Table 12: This table eontains variables that depend on pressure, voidage and temperature. Average values of these variables are used in simulation of the opera ti on of the multitubular reactor. The pressure in the multitubular reactor is atmospheric (1.01325 bar). The values in this tab Ie are used to roughly estimate the performance ofthe reactor T u _Pg Sc _DER kg [K) [mis) [kg/m3_{] [ -] [m}2_{/s1 [mis]} 625 3.09 0.57 2.03 0.00060 0.02402 650 3.22 0.55 2.02 0.00063 0.02529 675 3.34 0.53 2.01 0.00065 0.02656 700 3.47 0.51 2.01 0.00068 0.02785 725 3.59 0.49 2.00 0.00070 0.02914 750 3.71 0.48 2.00 0.00073 0.03045 average 3.40 0.523 2.01 0.000665 0.02720 21