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Chemical Engineering and Processing 46 (2007) 918–923

Ammoxidation of propylene to acrylonitrile in abench-scale circulating fluidized bed reactor

Yongqi Hu a,b,∗, Fengyun Zhao b, Fei Wei a, Yong Jin a

a Department of Chemical Engineering, Tsinghua University, Beijing 100084, Chinab Institute of Chemical and Pharmaceutical Engineering, Hebei University of Science

and Technolgy, Shijiazhuang 050018, China

Received 16 March 2007; received in revised form 21 May 2007; accepted 22 May 2007Available online 29 May 2007

bstract

The ammoxidation of propylene to acrylonitrile over Mo-Bi/�-Al2O3 catalyst was investigated in a bench-scale hot model riser reactor withmm i.d. and 30 m in length. Propylene conversion and product yields were investigated under various operation conditions and the optimum

onditions have been found for the new type reactor. The results show that the efficiency of catalyst is increased by four times and the yield ofcrylonitrile is increased by 3% for type A catalyst and by 6.5% for type B catalyst in comparison with a commercial turbulent fluidized bed reactor.he yield of acrylonitrile can be further increased through staged air feeding strategy.2007 Elsevier B.V. All rights reserved.

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eywords: Propylene; Ammoxidation; Acrylonitrile; Circulating fluidized bed

. Introduction

The heterogeneous selective ammoxidation of propylene intocrylonitrile (AN) is one of the most commercially significanteactions.

H3CH CH2 + NH3 + 32 O2

713−723 K,Catalyst−→ CH2 CHCN + 3H2O (1)

he features of this reaction include that: (1) it is a highlyxothermal reaction, �H = −512.5 kJ/mol; (2) the desired prod-ct acrylonitrile is a intermediate which may further be oxidizednto CO2 or CO, gas backmixing in reactor will cause the overox-dation of AN and thus, the decrease of AN yield; (3) it follows

edox mechanism, i.e., oxygen is supplied by catalyst in theorm of lattice oxygen and subsequently the reduced catalyst iseoxidized (regenerated) by molecular oxygen [1].

∗ Corresponding author at: Institute of Chemical and Pharmaceutical Engi-eering, Hebei University of Science and Technolgy, Shijiazhuang 050018,hina. Tel.: +86 311 88632175; fax: +86 311 88632175.

E-mail addresses: huyq@hebust.edu.cn, yongqi h@yahoo.com.cn (Y. Hu).

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255-2701/$ – see front matter © 2007 Elsevier B.V. All rights reserved.oi:10.1016/j.cep.2007.05.009

r reactor

Turbulent fluidized bed (TFB) reactor has been employedor propylene ammoxidation to synthesize AN for decades.ffective heat and mass transfer in TFB makes it advantageousver packed bed reactor on the control of reaction temperature.owever, TFB still suffers from severe axial gas and solids back-ixing, insufficient gas and solids contact and small throughput.oreover, it is difficult to build a catalyst regeneration region toeet the requirement of redox reaction mechanism in TFB due

o highly backmixing of gas and solids.To solve these problems, obstacles such as shaped metal-

ic articles, screens, grids, perforated plates, horizontal plates,ipes or the likes were laid in a catalyst bed to prevent theoalescence or growth of bubbles, or to prevent the back mix-ng of gas, thereby improving the contact between the feedas and the catalyst particles [2–6]. However these methodsre not practical because the construction for laying the obsta-les is complicated, and the mixing of the catalyst particless prevented by the obstacles and the distribution of the cat-lyst in the reactor becomes uneven in terms of space andime, so that it is difficult to stably and continuously con-

uct the operation [7]. A loop fluidized bed reactor with baffleor propylene ammoxidation was proposed and experimentallyxamined by Chen et al. [8]. A two stage fluidized bed was devel-ped for improving gas–solid contact, which can be applied in

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edox catalytic reaction such as the ammoxidation of propylene9].

The disadvantages of TFB can be overcome in a high-densityirculating fluidized bed (CFB) riser reactor [7,10–12]. Circu-ating fluidized bed allows the spatial separation of propylenemmoxidation in the riser reactor and catalyst regeneration inhe downcomer to maintain catalyst in oxygen-rich state for fur-her ammoxidation reaction. CFB riser operates under severalimes higher gas velocity than TFB, which decreases gas back-

ixing significantly. Staged addition of air can be permittedn CFB to control oxygen concentration along riser for opti-

al performance [12]. In conventional fluid catalytic crackingFCC), in which CFB is employed, the density of catalyst beds relatively low, however, high density in riser reactor, aver-ge catalyst fraction higher than 10% [13–15], is needed for theeaction of ammoxidation of propylene to acrylonitrile [7,10]or which longer gas–solid contact time is required than that forCC. Moreover, high-density operation allows higher mass andeat transfer to guarantee the conversion under higher operatingas velocity, and smaller in reactor size resulting in the decreasen construction cost [7].

This paper reports hot-model experiments on the selectivemmoxidation of propylene to acrylonitrile over Mo-Bi/�-l2O3 in a CFB riser reactor under high-density condition.onversions and product yields obtained in the reactor are com-ared with those measured in a commercial turbulent fluidizeded.

. Experimental

Mo-Bi/�-Al2O3 catalyst was used in the hot-model exper-ments. In order to obtain representative experimental results,he catalyst was taken from a commercial TFB reactor. Aftermployed for several months, the catalyst had been reaching itstable state of activity. The properties of the catalyst are listedn Table 1.

The bench-scale circulating fluidized bed reactor is shown inig. 1. It consisted of a riser reactor, a separator, a regenerator inhich catalyst was reoxidized by air, and an electrical heatinguidized bed bath. The riser, 0.007 m i.d. and 30 m in length,as spiraled round the regenerator. The long riser can simulta-eously guarantee a higher gas velocity (>3 m/s) and enough gasesidence time. The riser in spiraling type installed in a fluidized

eating bath, in order to maintain the isothermal conditions ofhe 30 m riser reactor. The inclination of the spiral tube againsthe horizontal was about 5◦. The temperature fluctuation of theuidized heating bath was controlled within 1 K. There was a

able 1he physical properties of the Bi/Mo catalyst

article size distribution (%)<45 �m 29.2>45 �m and <90 �m 62.3>90 �m 8.5

ensity of particle (kg/m3) 1800pecific surface area (m2/g) 0.68

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ig. 1. The schematic of the laboratory-scale high-density circulating fluidized-ed reactor.

ide air inlet on the riser at 10 m from the entrance of feed toxamine the staged air feed.

The flow rates of propylene, air and ammonia were controlledy mass flow controllers. The pipe of reactant air was placed inhe fluidized bed bath for preheating the reactant air, and thenhe air entered the bottom of the regenerator to fluidize the cat-lyst particles which flowed down to an injector. In the injector,eactant air and catalyst particles are mixed with propylene andmmonia up to the riser. The amount of carried catalyst wasontrolled by the flow rate of secondary air at the bottom of thenjector. In the riser the ammoxidation of propylene occurred,hen the catalyst particles were separated from the gas in theeparator and stored in a catalyst-collector. The catalyst parti-les were returned to the regenerator in batch to be reoxidizedy the regenerating air. The gaseous products from the separatorent to a combustor for venting or to an absorbing system for

nalysis.The residence time distribution of gas in the riser reactor was

easured by a pulse response method using a thermal conductiv-ty detector. The measured residence time distribution indicatedhat axial Peclet number (Pe) increased with gas velocity andas larger than 1000 when gas velocity was higher than 2 m/s,

ndicating that gas flow in the riser approached plug flow. Theverage catalyst volume fraction in the reactor was determinedy weighing the catalyst in the riser after suddenly closing theas feed. When gas velocity was 2.5–3.0 m/s and pressure dropas set to 0.03 MPa, the typical average catalyst fraction was.10–0.12, and the density of catalyst bed was 180–216 kg/m3,hich was 5–10 times higher than that of FCC riser. Nakamura

3

t al. [7] gave the density of catalyst bed was 100 kg/m orore, and preferable 200 kg/m3 or more for the ammoxidation of

ropylene in a circulating fluidized bed reactor. Although there isworld of difference on the structure between the above experi-

9 ng and Processing 46 (2007) 918–923

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ental riser reactor and commercial scale riser, the similarity onhe flow regimes, gas velocity and average catalyst fraction givesupport to the demonstration on the effectiveness to improvehe yield of AN in fast fluidizing flow regime compared to inurbulent fluidizing flow regime.

The gaseous products were collected in three 400 ml 0.1NNO3 scrubbers at 273 K. A temperature programmed FID gas

hromatograph was used to analyze AN, ACL, ACN and ACA.he gaseous products were analyzed by a TCD gas chromato-raph. Yield of HCN was determined by the addition of NaOH,ollowed by titration with 0.01 M AgNO3. Ammonia break-hrough was measured by titration of the HNO3 scrubber with.1N NaOH.

. Results and discussions

In order to demonstrate the features of CFB riser reactor forhe ammoxidation of propylene to AN, hot model experimentsere made under different operation conditions: contact time,

emperature, and feed ratio. The experiments with side feed ofir were also carried out. The results are discussed as follows.

.1. Effects of contact time on the product distribution

Product yield distributions are shown in Fig. 2 as a functionf the contact time, W/F. Steep increases in both propylene con-ersion and AN yield are observed in the beginning stage ofeaction. A complete conversion is nearly reached at the contactime longer than 125 g cat. h/mol C3H6. WWH, the weight (kg)f reacted propylene per kilogram catalyst per hour, is usually

sed to represent the efficiency of catalyst in a certain reac-or. For the catalyst used in the experiments, WWH is 0.065 inommercial TFB reactors. The contact time of 125 g cat. h/mol3H6 is equivalent to 0.33 in WWH, indicating that the capac-

ig. 2. Changes in yield of products with contact time: P = 0.10 MPa, T = 718 K,ir/C3 = 10.5, NH3/C3 = 1.15.

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ig. 3. Changes in yield of AN with temperature: P = 0.10 MPa, air/C3 = 10.5,H3/C3 = 1.15, W/F = 125 g h/mol.

ty of the catalyst in CFB is about four times higher than that inommercial TFBs.

As can be seen in Fig. 2, increasing contact time, AN yieldncreases significantly at beginning reaches a maximum and thenecreases gradually. The yields of both COx and HCN, however,ncrease continuously with contact time, indicating that AN canurther be oxidized to HCN and COx under long contact time.N is intermediate product, severe gas backmixing decrease AN

electivity and yield.

.2. Effect of reaction temperature

Fig. 3 plots the effect of temperature on AN yield. The high-st yield can be achieved within the range of 708–718 K. Underower temperature, AN yield is low due to the slow reactionate and thus, the low conversion level; under higher temper-ture, overoxidation causes the decrease of AN yield. On thether hand, COx production steadily increases as the reactionemperature increases. HCN yield changes insignificantly withhe increase of temperature.

.3. Effect of feed ratio

Air/C3 is an important controlling factor in industrial pro-esses. Theoretically, air/C3 ratio of 7.5 is enough for the maineaction to produce AN, in which 1.5 mol O2 reacts with 1 molropylene. Due to the existence of side reaction, air/C3 ratios usually maintained between 10 and 10.5 in commercial TFBeactors. Fig. 4 presents the effect of air/C3 on AN yield. Anptimum ratio of 9.5–10 is found. It is slightly lower than that in

he industrial process. This is partly because the regenerated cat-lyst carried some amount of oxygen in lattice type into the riser.ess overoxidation of AN resulted also in a low consumption ofxygen.

Y. Hu et al. / Chemical Engineering and Processing 46 (2007) 918–923 921

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ig. 4. Changes in yield of AN with air/C3 : P = 0.10 MPa, T = 718 K,H3/C3 = 1.15, W/F = 125 g h/mol.

NH3/C3 ratio is also important in the synthesis of AN. MoreH3 in the reacting gas promotes N-containing products, espe-

ially AN, and impedes O-containing products like COx andCL. The monotonous increase of AN yield and decrease ofOx yield can be seen in Fig. 5. NH3/C3 ratio of 1.1–1.2 isreferred. High NH3/C3 ratio will increase the operation costn the product separation.

.4. Comparison with commercial TFB reactor

Two type of Mo-Bi/Al2O3 catalysts was used in thexperiments. The experimental results from the experimentaligh-density riser reactor and the data from commercial TFBeactors are shown in Table 2. The results from a small bubble

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able 2omparison with commercial TFB reactor

Catalyst (type A)

HDCFB (reactor) TFB (reactor)

(K) 716 718(MPa) 0.081 0.05ir/C3 9.52 10.5H3/C3 1.19 1.15(m/s) 2.29 0.50

ield (%)AN 83.35 80.33ACL 0.31 0.1ACN 2.79 3.08HCN 5.66 5.91COx 6.98 10.21

99.09 99.63alance in C 1.04 –alance in O 0.99 –WH 0.349 0.065

ig. 5. Change in yield of AN with NH3/C3 : P = 0.10 MPa, T = 718 K,ir/C3 = 10.5, W/F = 125 g h/mol.

uidized bed (BFB) reactor for the evaluation of type A cat-lyst activity, provided by catalyst manufacturer, is also givenn this table. The experimental results indicate that comparedo the commercial TFB reactors, the high-density riser reactoras the following features (1) the operating gas velocity reaches.2–3 m/s, the throughput of reactants is increased by more thanour times; (2) the efficiency of the catalysts, WWH, is increasedy four times; (3) AN yield is increased by 3% for the type A cat-lyst and by 6.5% for the used type B catalyst, and the productionf COx are obviously decreased. The increases in reactor capac-

ty and in WWH are mainly due to better gas–solid contact iniser under high-density condition, fully regeneration of catalystn regenerator and the higher concentration of reactant gases inhe inlet of the riser reactor. The increase in AN yield is the result

Catalyst (used type B)

BFB (reactor) HDCFB (reactor) TFB (reactor)

718 724 7220.05 0.05 0.05– 9.7 10.08– 1.12 1.030.01 2.99 0.51

76.8 75.2 68.68– 0.0 0.79– 1.59 2.48– 10.98 4.21– 11.51 20.79

– 99.28 96.95– 1.08 –– 0.98 –– 0.451 0.065

922 Y. Hu et al. / Chemical Engineering an

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Fig. 6. Variations of AN yield with the fraction of oxygen side feed.

f nearly plug flow in the riser under higher gas velocity than thatn TFB.

.5. Experiments on staged air feeding

Essential plug flow and thus the existence of gradient of reac-ant concentration along riser make it possible to optimize theeed policy and to get high selectivity of the desired product.ow concentration of oxygen along riser will reduce oxida-

ion and favor ammoxidation, and thus will increase AN yield.owever, considering that too low O2/C3 ratio will cause over-

eduction of catalyst to lost its activity, only 10–30% of totalir was introduced through one side inlet at 10 m of the riser inhe experiments. In order to maintain similar state of flow in thexperiments under different side feed ratio, the air from side inletas replaced by 79% N2 + 21% O2 and only O2 was introduced

rom the side inlet while N2 entered the riser at the bottom of theiser. Fig. 6 shows the results of side oxygen feed experiments.

hen the experiments on side feed were carried out, the catalystad been run for several months and had experienced too manyimes of the rise and drop in temperature due to the start andtop of experiment. The catalyst activity was not so good as athe beginning, and AN yield under no side feed condition wasnly about 75%, as shown in Fig. 6. Nevertheless, the increasef nearly 5% in AN yield was obtained when 30% oxygen as theide feed, confirming the effectiveness of staged oxygen feedingtrategy.

. Conclusions

Hot model experiments were done on a laboratory-scale high-ensity CFB reactor with two kinds of Mo-Bi/�-Al2O3 catalystnder various operation conditions. The optimum operation con-

d Processing 46 (2007) 918–923

itions for high-density riser reactor are temperature 708–718 K,ir/propylene ratio 9.5, NH3/propylene ratio 1.1–1.2 and contactime 125 g h/mol C3H6. Compared with the commercial turbu-ent fluidized bed reactor, a high-density riser reactor has thedvantages including that (1) gas velocity reaches 2.2–3 m/s,nd the throughput of CFB reactor is increased more than fourimes; (2) the efficiency of catalyst, WWH, is increased by fourimes; (3) The yields of AN is increased by 3% for type A catalystnd by 6.5% for type B catalyst, and the yield of COx is obvi-usly decreased. Staged oxygen feeding can further promote thencrease of AN yield.

cknowledgements

Financial support from Petrochemical Company of China isratefully acknowledged. The authors are also grateful to Pro-essor Zhanwen Wang and Professor Zhiqing YU for their veryseful discussion and to Mr. Hongwei Dian, Xiaotao Wan andanhui Yang for their help on the experiments.

ppendix A. Nomenclature

CA acrylic acidCL acroleinCN acetonitrileir/C3 mole ratio of air and C3H6 (dimensionless)N acrylonitrileOx CO2 + COH3/C3 mole ratio of NH3 and C3H6 (dimensionless)

pressure at the exit of riser (MPa)reaction temperature (K)gas velocity in riser (m/s)

/F contact time on the basis of C3H6, g cat. h/mol C3H6WH the weight (kg) of reacted propylene per kilogram cat-

alyst per hourconversion ratio of propylene

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