A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance

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A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance Ourmazd Dehghani Khold, Mahboubeh Parhoudeh, Mohammad Reza Rahimpour∗, Sona Raeissi Department of Chemical Engineering, School of Chemical and Petroleum Engineering, Shiraz University, Shiraz 71345, Iran

a r t i c l e

i n f o

Article history: Received 7 September 2015 Revised 14 February 2016 Accepted 20 April 2016 Available online xxx Keywords: Acetylene hydrogenation Ethylene production Catalyst regeneration Tail-end reactor

a b s t r a c t Due to importance of acetylene removal from ethylene rich stream in polymerization units in order to prevent formation of highly explosive compounds, presenting any strategies improving the process efficiency has attracted researchers and industries interests. Proposing novel configurations in the reactionregeneration cycles with lower energy consumption and enhanced catalyst lifetime leading to profits for the polymerization units is of a great interest. In this regard, a new configuration for the tail-end acetylene hydrogenation reactor of a commercial olefin plant was proposed in the present study. Based on the proposed configuration, a new protocol is applied to the process for utilizing total capacity of active catalysts in each fixed bed. A mathematical model based on component and energy balance was developed and validated to evaluate the operability of the proposed configuration. Operation of different runs in the new protocol was compared with the conventional configuration. The obtained results revealed that the new configuration is managed to decrease the volume of the regenerated catalyst in a 4-year period of operation. A 33.3% decrease in the volume of the regenerated catalyst was observed. Such an achievement was obtained without reducing acetylene conversion and ethylene yield. © 2016 Taiwan Institute of Chemical Engineers. Published by Elsevier B.V. All rights reserved.

1. Introduction 1.1. Ethylene and acetylene Ethylene, which is a colorless, nonpolar and highly flammable gas, could be used as a raw material for diverse organic compounds. Polyethylene, polyvinyl chloride (PVC), ethyl benzene and polyesters are generally produced from ethylene and its derivatives due to low production costs. Besides, its reaction with other low cost accessible material such as oxygen and water leads to production of beneficial chemicals. The major application of ethylene monomers is in polymerization units in which ethylene is converted to polyethylene [1]. Ethylene is mainly produced by thermal cracking of various feedstocks such as ethane and naphtha. However propane, butane and gas oil are other alternatives for ethylene production. Ethane and naphtha cracking for ethylene production is carried out in gas and liquid furnaces, respectively. Due to increasing growth in ethylene supply and demand, any enhancement in the process of ethylene production may lead to remarkable profits for the petrochemical industries [2,3]. Acetylene is an undesired side product in ethylene manufacturing process. Depending on the feed and cracking conditions, ∗

Corresponding author. Tel.:+98 711 2303071; fax: +98 711 6287294. E-mail address: [email protected] (M.R. Rahimpour).

about 0.5–2.5 ton acetylene is produced per 100 ton of ethylene. Acetylene is a contamination for catalyst of the polymerization unit. Maximum acetylene content for catalyst of the polymerization cannot exceed 1 ppm; otherwise, it leads to formation of metal acetylides, which are highly explosive. Hence, acetylene has to be eliminated from ethylene to assure an acetylene free stream for the polymerization unit [4,5]. An olefin plant in a domestic petrochemical complex uses thermal cracking of alkanes such as ethane, propane, butane, naphtha and gas oil to produce 1324,0 0 0 ton of ethylene per year [6]. Both gas and liquid crackers are applied to produce ethylene using ethane and naphtha as a feedstock, respectively [6–8]. Typical analysis of both liquid and gas furnace output streams in this plant are shown in Table 1. As clear, ethylene percentage in output stream of the gas furnace is considerably higher than that of the liquid furnace. Besides, more acetylene is produced in the liquid furnace. 1.2. Hydrogenation techniques Generally, acetylene is eliminated through two different strategies, removal from the main stream and hydrogenation. The former is not cost effective; therefor it is unfavorable for large-scale applications. Conventionally, catalytic hydrogenation in an adiabatic fixed bed reactor is running the art of acetylene elimination from ethylene. Defects in the performance of an acetylene hydrogenation reactor may lead to remarkable extra charges [5,7–10].

http://dx.doi.org/10.1016/j.jtice.2016.04.027 1876-1070/© 2016 Taiwan Institute of Chemical Engineers. Published by Elsevier B.V. All rights reserved.

Please cite this article as: O. Dehghani Khold et al., A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance, Journal of the Taiwan Institute of Chemical Engineers (2016), http://dx.doi.org/10.1016/j.jtice.2016.04.027

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Nomenclature a AC Cpg Ct CR Ea,i F Ft H kd ki NP Pi ri

catalyst activity [dimensionless] cross section [m2 ] specific heat of the gas at constant pressure [J/mol K] total concentration [mol/m3 ] crossover probability constant [dimensionless] activation energy [J/mol] DE-step size [dimensionless] total molar flow rate [mol/s] hydrogen [dimensionless] deactivation rate constant [day−1 ] reaction rate constant [bar−1 ] or [kmol/kgCat s bar2 ] number of population [dimensionless] partial pressure of component i [Pa] rate of reaction of component i [mol/kgCat s]

Table 1 A typical analysis of both liquid and gas furnace output streams in the domestic petrochemical planta . Component

Gas furnaces (wt%)

Liquid furnaces (wt%)

Acetylene Ethylene

0.34 50.75

0.86 34.11

a Sample method: gas chromatography ASTM D2504 test method for noncondensable gases in C2 and lighter hydrocarbon products and ASTM D6159 for other components by gas chromatography in olefin plant laboratory.

t T yi z

time [s] temperature [K] mole fraction of component i [dimensionless] axial reactor coordinate [m]

Greek letters H i enthalpy of formation of component i [J/mol] εB bed porosity [dimensionless] η catalyst effectiveness factor [dimensionless] ρB density of catalytic bed [kg/m3 ] Subscripts 1 bed 1 (lead bed) 2 bed 2 (guard bed) i numerator for component/day In inlet j numerator for reaction Out outlet

There are diverse configurations for the acetylene hydrogenation reactor in an industrial olefin plant, including; raw gas catalytic reactor, front-end reactor, and tail-end reactor, among which, tail-end and front-end configurations cover 64% and 29% of all operating plants, respectively [5]. A schematic diagram of an industrial olefin plant is sketched in Fig. 1. As shown, the acetylene hydrogenation reactor could be devised in three different locations in each configuration. The specified location 1 in Fig. 1 shows the raw gas catalytic hydrogenation reactor suggested for acetylene hydrogenation in the

Fig. 1. Schematic diagram of an olefin plant.

Please cite this article as: O. Dehghani Khold et al., A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance, Journal of the Taiwan Institute of Chemical Engineers (2016), http://dx.doi.org/10.1016/j.jtice.2016.04.027

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presence of considerable sulfur quantity in the inlet feed of the plant. In this configuration, the effluent cracked gas is compressed and sent to the hydrogenation reactor. The feed of this reactor includes all types of hydrocarbon groups and sulfur compounds. Hence, as palladium and platinum catalysts can easily be poisoned by sulfur compounds, these active metals must not be applied in this configuration [11]. The specified location 2 in Fig. 1 shows the front-end catalytic hydrogenation reactor. In such a configuration, sulfur compounds are primarily omitted from the streams using the installed caustic tower. Thereupon, active metals such as palladium can be employed as a catalyst to enhance the selectivity [12]. In this system, the feed stream often contains a great amount of hydrogen (10– 35 mol%) leading to diminishing the coke deposition [9,13]. Gobbo et al. investigated modeling, simulation, and dynamic optimization of this sort of acetylene hydrogenation reactor [9]. The most common configuration for acetylene hydrogenation in an olefin plant is tail-end. The specified location 3 in Fig. 1 shows such a configuration. Lighter gases such as CH4 and CO are removed from the ethylene stream. The ethylene rich stream containing ethylene, acetylene and ethane goes through the acetylene hydrogenation reactor. Besides, additional hydrogen is injected into the system [14]. More coke deposition in the reactor and effluent streams contaminated with hydrogen and methane are drawbacks of this configuration [15]. Amount of the acetylene impurity in the ethylene stream is a determinant factor in selecting number of the utilized reactors. Accordingly, for ethylene streams with acetylene content lower than 50 0 0 ppm, a single hydrogenation reactor is utilized. Two reactors are required to omit acetylene from streams with 0.5–1.7 mol% acetylene, and for concentrations more than 1.7 mol%, three reactors are considered [8]. 1.3. Hydrogenation catalysis Catalytic hydrogenation reaction was first brought up in a study by Anderson et al. in 1948 in which they used nickel sulfide catalyst based on alumina or silica for hydrogenation of diolefins to monoolefins [16]. Catalysts made of calcium, barium, cadmium, strontium as well as magnesium on Cr2 O3 [17], zinc and nickel on alumina and silica [18,19] and copper on silica, alumina and magnesia [20] were tested for acetylene hydrogenation in different hydrocarbon streams. Palladium deposited on the alumina was then found to have the most selectivity and resistance for acetylene hydrogenation among different catalysts [21–23]. Godinez et al. investigated front-end and tail-end configurations for selective hydrogenation of C2 –C3 mixtures on two types of promoted Pd/Al2 O3 catalysts [15,21]. About 0.015–0.05 wt% of palladium beside metallic promoters such as silver is commercially used for hydrogenation of acetylene [14,21,24–26]. 1.4. Catalyst deactivation Chemical or physical phenomena occurring simultaneously with the main reactions make the activity of catalysts change. On the other hand, catalyst deactivation is the overall result of the decline in active sites of the catalytic surface, which is caused by a variety of mechanisms, both chemical and physical [27]. In hydrogenation of acetylene, catalyst performance declines gradually due to coke deposition. Oligomer (green oil) consisting unsaturated aliphatic hydrocarbons are the main cause of the catalyst deactivation in hydrogenation of acetylene [5,28]. Green oil precipitation creates a sticky layer on the Pd/Ag/α -Al2 O3 catalyst surfaces preventing small acetylene molecules to diffuse through the pores. Such an inhibition effect leads to reducing accessible active sites. As a result, catalyst activity and ethylene selectivity declines remarkably. Fresh and deactivated catalyst samples of the acetylene hydro-

3

Fig. 2. Fresh and coked sample of Pd–Ag/α -Al2 O3 catalysts. Table 2 Specifications of the catalyst samples in hydrogenation unit of the domestic petrochemical plant. Properties

Values

Size and shape Density of catalyst Carrier SA Pd content Ag/Pd ratio CO require Ext. surface of particle (SP ) Volume of particle (VP ) Heat capacity of catalyst (CS ) Thermal conductivity of catalyst (KS ) Particle diameter (dp ) Particle porosity (ε p ) Particle tortuosity Average pore diameter Specific internal BET ө-Al2 O3 content of carriera α -Al2 O3 content of carrier γ -Al2 O3 content of carrier

2–4 (mm) and sphere 1400 (kg/m3 ) Medium 300 (ppm) with support Al2 O3 6/1 0.5–2 (ppm) optional 2.827 × 10−5 (m2 ) 1.414 × 10−3 (m3 ) 10 0 0 (J/kg K) 0.28211 (W/m k) 4 (mm) 60% 2.5 30 (nm) 40 (m2 /g) 72% 16% 12%

a Crystal structure of the carrier (Al2 O3 ) determine by XRD (X-ray diffraction) test and method of BS EN 13925-1.

genation unit in the domestic petrochemical complex are shown in Fig. 2. Catalyst specifications are presented in Table 2 [8]. Due to presence of valuable and expensive metals such as silver in the catalysts, frequent replacement of fresh catalysts is not cost effective (typically, Pd/Ag/α -Al2 O3 catalyst costs 82.5 EUR/kg based on technical bulletin of C2 tail-end acetylene converters [6]). Thus, search for efficient methods of catalyst regeneration cycles is economically of a great interest. 1.5. Catalyst regeneration Deactivated catalyst by coke deposition must be periodically regenerated. Conventionally, thermal regeneration at 40 0–80 0 °C is employed. For this purpose, fixed bed reactors are shut down, an inert gas is passed through the bed to remove the remaining reacting mixture and then oxygen or air stream is passed through the bed in order to burn off the coke. As a result, an exothermic process occurs during coke burning, the reaction zone moves through the bed, the generated heat is transported by the flowing gas. However, the catalyst particles may temporarily retain part of the generated heat [29,30]. Due to highly exothermic nature of decoking reactions, high temperature may cause sintering of the catalyst particles and physical damages to the process equipment [31,32]. Hence, the regeneration step is industrially vital. Besides, increasing the number of times the catalysts are thermally regenerated may be accompanied by irreversible changes in the internal structure of the catalyst as well as raising energy consumption [30]. Elegant configurations in the reaction-regeneration

Please cite this article as: O. Dehghani Khold et al., A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance, Journal of the Taiwan Institute of Chemical Engineers (2016), http://dx.doi.org/10.1016/j.jtice.2016.04.027

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Fig. 3. Acetylene hydrogenation reactor in run 1 of the conventional configuration.

cycles with lower energy consumption lead to increase catalyst lifetime and efficiency and consequently process performance. Thus, any enhancement in the reaction-regeneration cycles may lead to remarkable profits for the industries. In the previous works published from our group, Rahimpour et al. and Dehghani et al. proposed new configurations for decoking of acetylene hydrogenation catalyst. They suggested separate regeneration of each bed to assure complete coke burning in less regeneration time [5,33]. 1.6. Objective In the present study, a new configuration for reactionregeneration cycle in hydrogenation of acetylene in a commercial olefin plant is proposed. The main idea of the proposed configuration is to arrange reactors in a manner to utilize total capacity of catalyst filled in each bed. A mathematical model based on mass and energy balance in the presence of catalyst deactivation is proposed. Operability of the proposed configuration against the conventional one is compared in terms of acetylene removal, ethylene yield and volume of the regenerated catalyst in certain period. 2. Reaction-regeneration cycle of tail-end acetylene hydrogenation 2.1. Conventional tail-end configuration The tail-end acetylene hydrogenation reactors of the domestic petrochemical plant consists of two fixed beds with inter-stage cooling as shown in Figs. 3 and 4. The first bed called lead bed (LB) is utilized for the elimination of 60–70% of acetylene. A rapid rise in temperature of the lead bed is observed due to highly exothermic nature of acetylene hydrogenation reaction, which may cause catalyst sintering. Hence, total elimination of acetylene is not possible in one bed. Therefore, remaining acetylene is eliminated in the second bed called guard bed (GB). Utilizing the guard bed always keeps the acetylene concentration below the desired level i.e. 1 ppm.

As shown in Fig. 3, pure hydrogen is added to the ethylene rich stream of de-ethanizer and the mixture is heated up to 35–40 °C, which is often identified as the Start-Of-Run (SOR) temperature. Then, it is sent to LB. The outlet stream of LB is cooled to the SOR temperature of GB and then mixed with the pure hydrogen. The outlet stream of GB is subsequently cooled down and injected into a scrubber called green oil knock-out pot. High concentration of acetylene in the initial feed may produce a large amount of C4+ and green oil. These chemicals might interfere with the operation of the C2 fractionators and as a consequence should be eliminated by green oil knock-out pot. While, beds 1 and 2 (i.e. lead bed and guard bed) are utilized, bed 3 and 4 are out of service. After a 360-day operation, catalysts of LB are entirely deactivated. Then, both LB and GB are sent to the regeneration unit simultaneously (see Fig. 4). However, catalysts of GB are not totally deactivated. Subsequently for next run, as depicted in Fig. 4, the hydrocarbon feed is injected to the second reactor (spare reactor containing fresh LB and GB). Operating parameters and characteristics of the acetylene hydrogenation reactor and compositions of the inlet stream under normal operating condition (i.e., the number of gas and liquid furnaces are the same under the normal operating condition) are tabulated in Tables 3 and 4, respectively. Besides, a typical output analysis of the acetylene hydrogenation reactor is tabulated in Table 5. As clear, the quantities of side products (such as butane, butene, butadiene, propane, propylene, and C6 +) are small that they could be neglected. 2.2. Problem detection As previously mentioned, about 70% of inlet acetylene is eliminated in LB and the remaining acetylene converts in GB. An equal amount of fresh catalyst is loaded into LB and GB. Hence, the catalyst of LB is deactivated faster than that of GB because of exposure to higher amount of acetylene. After a 360-day operation, catalyst of LB is entirely deactivated while, considerable part of catalyst in GB is still active. In the conventional configuration (Fig. 4) both

Please cite this article as: O. Dehghani Khold et al., A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance, Journal of the Taiwan Institute of Chemical Engineers (2016), http://dx.doi.org/10.1016/j.jtice.2016.04.027

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5

Fig. 4. Acetylene hydrogenation reactor in run 2 of the conventional configuration.

Table 3 Operating parameters and characteristics of the acetylene hydrogenation reactor in the domestic petrochemical plant under normal operating condition. Parameters

Components

Quantity

Mass flow rate (kg/h)

Ethylene Ethane Hydrogen n-Butane i-Butane 1-Butene 1,3-Butadiene cis-2-Butene trans-2-Butene C6+ Acetylene Propylene Propane

69.84 mol% 29.90 mol% 11.3 ppm 0.0 0 04 mol% <0.0 0 01 mol% 0.0193 mol% <0.0 0 01 mol% 0.004 mol% 0.0085 mol% <0.0 0 01 mol% <1 ppm 0.1082 mol% 0.0012 mol%

136.6 × 103 62.68 × 103 – – – – – – – – – – –

Value

Operating conditions Pressure Temperature (start-of-run) Temperature (end-of-run) Total mass flow rate

25 (bar) 35 (°C) 69.5 (°C) 20 0,0 0 0 (kg/h)

Specifications of the reactor Bulk density (ρB ) Bed porosity (εB ) Reactor diameter (D) Length of bed (L) Effectiveness factor (η) Catalysts volume of bed

720 (kg/m3 ) 0.40 3.92 (m) 3.35 (m) 1 40 (m3 )

Table 4 A typical analysis of the inlet stream of the acetylene hydrogenation reactor. Components

mol%

wt%

Mass flow rate (kg/h)

Methane Acetylene Ethylene Ethane MAPDa Propylene Propane

0.02 0.92 69.329 29.56 0.0 0 02 0.17 0.001

0.0099 0.83 67.88 31.02 0.0 0 04 0.2481 0.0047

19.80 1.66 × 103 135.76 × 103 62.04 × 103 0.86 496.20 9.40

Sum

100

100

200 × 103

a

Table 5 A typical analysis of the outlet stream of the acetylene hydrogenation reactor.

MAPD: methyl acetylene propadiene.

totally deactivated LB and partly deactivated GB are sent to the regeneration step simultaneously. On the other hand, the active fraction of catalysts in GB is sent to regeneration step. Extra energy is consumed for fresh catalyst and regeneration time increases. Besides, it causes destruction and remarkably reduction in activity of the fresh catalysts after regeneration process. In other word, in this

manner, energy and time are wasted and it may cause some damages to active catalyst particles, which imposes extra charges on the process. 2.3. The new configuration In order to solve the aforementioned problem in the conventional configuration of acetylene hydrogenation reactor, a new configuration is suggested. A protocol with different runs is suggested in the proposed configuration, as presented in Table 6. The first run is similar to the conventional configuration. After a 360-day operation, catalysts of bed 1 are totally deactivated. In run 2, the fresh feed from de-ethanizer is injected to the bed 2 (now working as a lead bed) containing 40–50% active catalyst and the outlet stream of this bed is sent into the bed 3 (now working as a guard bed) containing fresh catalyst (see Fig. 5(a)). In this manner, the whole catalysts are utilized in the process and energy and time are not spent on the regeneration of active catalyst. Hence, the remaining active catalysts in the bed 2 deactivate totally in run 2 and after

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Table 6 Sequence of beds swapping for the conventional and proposed configurations. Run no.

Run 1

Proposed configuration Lead bed Bed 1 Guard bed Bed 2 Actual configuration Lead bed Bed 1 Guard bed Bed 2

Run 2

Bed 2 Bed 3 Bed 3 Bed 4

Run 3

Bed 3 Bed 4 Bed 1 Bed 2

Run 4

Bed 4 Bed 1 Bed 3 Bed 4

Table 7 Chemical reactions, corresponding heat of reactions and rate equations.

Hr

No. of reaction

Reaction

(1)

C2 H 2 + H 2 → C2 H 4

−42.2

rC2 H2 = (1+k P 2 )(2 1+2k P ) 2 C2 H2 3 H2

(2)

C2 H 4 + H 2 → C2 H 6

−33.1

rC2 H6 =

(3)

C2 H 2 + 2H 2 → C2 H 6

−75.3

Neglected

Rate equation

(kcal/mol) k1 PC

H

PH

k4 PC2 H4 PH2

(1+k5 PC2 H4 )1.25 (1+k6 PH2 )

Table 8 Constant parameters of the rate equations. Arrhenius equation

ki = ki,0 exp(−E a,i /RT )

Rate Rate constants

ki, 0 (kmol/kgCat s bar2 or bar1 )

Ea,i (kJ/kmol)

k1 k2 k3 k4 k5 k6

48.01 584.59 2.86 202.67 0.07 2.89

1220.5 −5558.7 −3361.3 39774.2 −12493.4 −3325.6

configuration, a mathematical model is developed in the following sections. 3. Reaction scheme and kinetic expressions Three dominant reactions occurring in acetylene hydrogenation process are acetylene hydrogenation to ethylene, ethylene hydrogenation to ethane and direct hydrogenation of acetylene to ethane [4,29–31]. The aforementioned reactions, corresponding heat of reactions and rate equations expressed in terms of components partial pressures are given in Table 7. Leviness et al. and Sarkani et al. found out that direct hydrogenation of acetylene to ethane could be ignored comparing hydrogenation of acetylene to ethylene and hydrogenation of ethylene to ethane. Therefore, reactions 1 and 2 in Table 7 are the predominant ones occurring in the systems [32,33]. Mostoufi et al. established a kinetic model for the reactions 1 and 2 based on experiments in a fixed bed reactor. By applying the Arrhenius equation, they also evaluated the constant parameters of the rate equations. The obtained constants are tabulated in Table 8 [21]. 4. Mathematical modeling 4.1. Governing equations Fig. 5. (a) Acetylene hydrogenation reactor in run 2 of the proposed configuration. (b) Acetylene hydrogenation reactor in run 3 of the proposed configuration.

completion of run 2, beds 1 and 2 are sent to regeneration unit simultaneously (see Fig. 5(b)). Subsequently, as depicted in Fig. 5(b), for run 3 the fresh feed from de-ethanizer is injected to the bed 3, which is partly deactivated in run 2. In run 3, bed 3 works as a lead bed and fresh bed 4 works as guard bed. This trend continues like a “merry-go-round” for all subsequent runs. The sequence of the beds shifting is displayed clearly in Table 6 for four successive runs in the proposed and conventional configurations of acetylene hydrogenation reactor. With this protocol, optimal use of total catalyst volume without extra energy and time consumption for regeneration of unused catalyst is managed. However, new equipment is not added to the process. On the other hand, these advantages are obtained with flow rearrangement. Hence, no extra charges are imposed on the process. To evaluate the operability of the proposed

To evaluate the operability of conventional and proposed configurations, a one-dimensional model based on the following assumptions is applied: • • •

•

• • • • • •

A homogeneous model is considered. The reactor operates at unsteady state condition. Radial variation in temperature and concentration is neglected (one-dimensional model). Axial heat and mass diffusion are negligible due to high gas velocity. Plug flow pattern is considered. Catalyst deactivation is considered. Bed porosity is constant leading to symmetry in the reactor. The gas mixture is considered as an ideal gas. Lateral heat loss is neglected (adiabatic reactor). Pressure drop along the reactor is negligible.

In order to obtain appropriate heat and mass balance equations, an axial differential element is selected. Consequently, a set of

Please cite this article as: O. Dehghani Khold et al., A new configuration in the tail-end acetylene hydrogenation reactor to enhance catalyst lifetime and performance, Journal of the Taiwan Institute of Chemical Engineers (2016), http://dx.doi.org/10.1016/j.jtice.2016.04.027

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O. Dehghani Khold et al. / Journal of the Taiwan Institute of Chemical Engineers 000 (2016) 1–14 Table 9 Obtained activity models for the conventional and proposed configurations based on some industrial data. Lead bed

tions and corresponding initial and boundary conditions [34]:

εB Ct

Guard bed

Proposed configuration Run 1 a (t ) = exp(−0.0053t ) Run 2 a (t ) = exp(−0.0017t ) − 0.5552 Run 3 a (t ) = exp(−0.003t ) − 0.2687

a (t ) = exp(−0.0022t ) a (t ) = exp(−0.0015t ) a (t ) = exp(−0.0019t )

Conventional configuration Run 1 a (t ) = exp(−0.0053t ) Run 2 a (t ) = exp(−0.0053t )

a (t ) = exp(−0.0022t ) a (t ) = exp(−0.0022t )

differential equations with axial direction length and time as independent variables is obtained. Followings are the governing equa-

7

Ft ∂ yi ∂ yi =− . + η ρB ri a ∂t Ac ∂ z

εB Ct Cpg

2  ∂T Ft ∂ T = − Cpg . + η ρB a r j ( − H j ) ∂t Ac ∂ z j=1

z=0

yi = yi, in

T = Tin

t=0

a = a (0 ) = 1

Tin = Tin (0 )

(1)

(2)

(3) (4)

where i represents the components in the reactions (hydrogen, ethane, ethylene and acetylene) and j denotes the reactions (reactions 1 and 2 in Table 7). Equation 1 is the material balance of the reacting bed with assumptions summarized above. The first term

Fig. 6. (a) Acetylene conversion in the lead bed (LB) of the conventional configuration. Comparison of model results with plant data during a 360-day period of operation. (b) Lead bed (LB) outlet temperature of the conventional configuration. Comparison of model results with plant data during a 360-day period of operation. (c) Guard bed (GB) outlet temperature of the conventional configuration. Comparison of model results with plant data during a 360-day period of operation.

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on the right hand side of equation 1 is bulk convection of the materials, while the second term is the consumption/generation rate of the components. Accumulation of the components is given by the term on the left hand side of Eq. (1). Eq. (2) is the energy balance of the reacting bed with the previously summarized assumptions. The first term on the right hand side of this equation is to consider bulk convection of the heat, while the second term is heat generated by the reactions. The term on the left hand side of this equation is to account for unsteady state condition. As a whole, a homogeneous ideal reacting system in a one-dimensional plug flow unsteady reactor is considered. As declared in the reactor configuration, Tin is the inlet temperature of the beds varying with time. The inlet conditions of guard beds are the same as outlet conditions of the lead beds. Auxiliary correlation to account temperature dependent of heat capacities of components is presented in Appendix A.

Table 10 Comparison between plant data of acetylene hydrogenation reactor and results of the model. Variable

%ARDa

Acetylene conversion Lead bed outlet temperature Guard bed outlet temperature Average

4.6 6.8 8.5 6.6

Variable

Plant

Output ethylene (kg/h) Output ethane (kg/h)

136.60 × 10 62.68 × 103

a

% ARD =

100 NDP

|

X ind. −X mod. X ind.

Model 3

%ARD

137.20 × 10 63.20 × 103 3

0.44 0.83

|.

4.2. Catalyst deactivation model As mentioned previously, catalyst performance declines gradually due to coke deposition. Considering a proper deactivation equation in the mathematical model could help to find better understanding from system behavior and the catalyst lifetime. Catalyst deactivation model for the commercial acetylene hydrogenation process adopted from Mostoufi et al. is as follows [21]:

da = −kd a dt

(5)

a(t ) = exp(−kd t )

(6)

where a and kd are activity of the catalyst and the deactivation constant, respectively. a and kd could be determined based on some industrial data. Based on some industrial data obtained in the conventional configuration, activity of bed 1 (i.e. lead bed) is about 15% after a 360-day period of operation. Besides, approximately 70% of all inlet acetylene is eliminated in the first bed and the remaining is eliminated in the second one. Accordingly, the ratio of deactivation constant of lead bed to guard bed is assumed 7 to 3. In further runs in the proposed configuration, deactivation constants and deactivation model were obtained based on the final deactivation of the previous runs. The obtained results are shown in Table 9. 4.3. Numerical solution A set of differential equations coupling with associated initial and boundary conditions are solved numerically to evaluate thermal and molar behavior of the proposed configurations. It should be mentioned that the reactor length is divided into 100 grids leading to a grid independent solution with low numerical error and good convergence and stability. 5. Results and analysis 5.1. Model validation The proposed model is validated against data of acetylene hydrogenation reactor of the domestic petrochemical plant under the design specifications and input data listed in Tables 3 and 4. The comparison between simulation results and the plant data for the conventional acetylene hydrogenation reactor is shown in Fig. 6. Clearly, it is observed that the predictions of model are in good agreement with the plant data and the proposed model performs well under special case of industrial conditions. Fig. 6(a) shows the model predictions and output plant data of acetylene conversion in LB in the conventional configuration. A significant decrease in acetylene conversion due to gradual decay in

Fig. 7. Relative errors of acetylene conversion and LB and GB temperature profiles.

catalyst activity is observed. Besides, a good agreement between plant data and the simulation results is observed clearly. The outlet temperatures of conventional LB and GB during a 360-day operation and corresponding simulation results are also compared in Fig. 6(b) and (c), respectively. Clearly, an acceptable agreement between the plant data and the simulated ones are observed. A quantitative comparison between the plant data and the simulation results in terms of average relative deviation (%ARD) is carried out in Table 10. As can be seen, %ARD for acetylene conversion and LB and GB temperature profiles are less than %9 and output mass flow rates of ethylene and ethane at the final day of operation are less than %1, concluded in acceptable performance of the proposed model under the industrial condition of conventional acetylene hydrogenation reactor. Besides, relative error of each data point for acetylene conversion and output temperatures are shown in Fig. 7. As can be seen, almost all the errors are less than 10% indicating acceptable performance of the mathematical model.

5.2. Conventional and new configurations In the following sections, to evaluate the operability of the conventional and proposed configurations, predictions of the mathematical model will be discussed in detail for runs 1–3.

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Fig. 8. (a) Catalyst activity profile for both lead bed and guard bed in the conventional configuration (i.e. run 1). (b) Acetylene molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in the conventional configuration (i.e. run 1). (c) Ethane molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in the conventional configuration (i.e. run 1). (d) Ethylene molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in the conventional configuration (i.e. run 1).

Obtained results for the conventional configuration (i.e. run 1) are shown in Fig. 8. The activity profile for both lead bed and guard bed is shown in Fig. 8(a). Higher deactivation rate in the lead bed is observed due to exposure to higher acetylene concentration. As clear, after a 360-day operation, catalyst of guard bed is still 40– 50% active. Hence, it could be utilized in run 2. Acetylene, ethane and ethylene molar flow rate versus dimensionless catalyst volume in two series catalyst beds in different days of operation are shown in Fig. 8(b)–(d), respectively. As can be seen in Fig. 8(b), although catalyst activity decays with time, total acetylene removal is managed utilizing two series catalyst beds. Run 2 in the proposed configuration is schematically shown Fig. 5(a). Obtained results for the run 2, in which partly active

bed 2 works as lead bed and completely fresh bed 3 works as guard bed are shown in Fig. 9. Based on activity profile shown in Fig. 9(a), process could operate according to run 2 arrangement in a 210-day period. After a 210-day operation, activity of bed 2 (i.e. lead bed in run 2) drops below 15% which means that run 2 is terminated. Acetylene, ethane and ethylene molar flow rates changes with catalyst volume in two series catalyst beds in different days of operation of run 2 are shown in Fig. 9(b)–(d), respectively. Clearly, acetylene is eliminated in bed 3 (i.e. guard bed in run 2). However, catalyst of lead bed is partly active in the beginning of run 2. After 210-day operation of run 2, 70–80% of bed 3 is still active. Hence, it could be utilized as lead bed in run 3. A schematic diagram of run 3 in the proposed configuration is shown in

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Fig. 9. (a) Catalyst activity profile for both lead bed and guard bed in run 2 of the proposed configuration. (b) Acetylene molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in run 2 of the proposed configuration. (c) Ethane molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in run 2 of the proposed configuration. (d) Ethylene molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in run 2 of the proposed configuration.

Fig. 5(b). As mentioned previously, catalyst bed 3 that is mainly active, works as lead bed and bed 4 that is fresh, works as guard bed. Besides, bed 1 and bed 2 are sent to the regeneration step. Results of numerical simulation for the run 3 are shown in Fig. 10. According to Fig. 10(a), run 3 could operate in a 315-day period. Then, activity of bed 3 (i.e. lead bed in run 3) drops below 15% and this is considered as the termination of run 3. Fig. 10(b)–(d) shows acetylene, ethane and ethylene molar flow rates changes with catalyst volume in lead and guard beds in different days of operation of run 3. Although, catalyst of lead bed in the beginning of run 3 is partly active, acetylene is eliminated in bed 4. At the end of 315day period, bed 4 is 50–60% active. Therefore, it could be utilized as lead bed in run 4.

In Fig. 11, acetylene, ethane and ethylene molar flow rates versus catalyst volume in lead and guard beds in the first and last days of operation in runs 1–3 are compared. As per Fig. 11(a), acetylene is totally removed in the guard beds in all of runs. However, activity of lead beds and deactivation rates are different in the beginning of runs. An important factor in the acetylene hydrogenation reactor is ethylene yield. On the other hand, it is desired to convert acetylene to ethylene rather than ethane. Ethane and ethylene profiles in runs 1–3 are shown in Fig. 11(b) and (c), respectively. Besides, Average overall ethylene yield in runs 1–3 are compared in Fig. 12. As can be seen, a slight increase in ethylene yield is obtained in runs 2 and 3.

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Fig. 10. (a) Catalyst activity profile for both lead bed and guard bed in run 3 of the proposed configuration. (b) Acetylene molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in run 3 of the proposed configuration. (c) Ethane molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in run 3 of the proposed configuration. (d) Ethylene molar flow rate versus dimensionless catalyst volume in two series fixed beds in different days of operation in run 3 of the proposed configuration.

Temperature profiles of lead and guard beds in runs 1–3 are compared in Fig. 13. Thermal pattern of the reactor is critical not only to the safety but also to the efficiency of the process. As clear, runs 1–3 have the same pattern in the first day of operation, i.e. a strictly increasing trend in the lead bed due to exothermic nature of the existing reactions and a smooth trend in the guard bed due the reached equilibrium in the reactions. The same pattern is observed in the last days of operation. However, the observed temperature variation of different runs at final days is less comparing those of earlier days. On the other hand, temperature change in runs 1–3 does not differ much at the end of operation. This is due to lower catalyst activity in the last days of operation. Catalyst activities of all the runs in the last days are almost the same but

catalyst activities in the first days are distinct. Hence, different reaction rates in the earlier days leads to discrepancies in the temperatures of diverse runs and almost the same reaction rates in the last days of operation leads to similar temperatures. Situation of each catalyst bed during a 4-year (1440 days) period of operation is determined in the conventional and new configurations in Table 11. Obviously, the proposed configuration requires less regeneration steps in 4 years. In addition to preventing regeneration of fresh catalysts, the regeneration process is performed two times instead of three times in 4 years. Volume of the regenerated catalysts and average overall ethylene yield in 4 years for two configurations is compared in Fig. 14. A 33.3% decrease in the volume of the regenerated catalyst is observed.

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Fig. 11. (a) Acetylene molar flow rates versus catalyst volume in lead and guard beds in the first and last days of operation in runs 1–3. (b) Ethane molar flow rates versus catalyst volume in lead and guard beds in the first and last days of operation in runs 1–3. (c) Ethylene molar flow rates versus catalyst volume in lead and guard beds in the first and last days of operation in runs 1–3.

Hence, in the proposed configuration, less catalysts degradation and higher catalysts lifetime is achieved. Besides, cost of regeneration is considerably reduced in this way. A slight increase in the average overall ethylene yield is also obtained. Indeed, total acetylene removal, a slight increase in ethylene yield, higher catalyst performance and less regeneration charges are obtained with utilization of the new configuration. 6. Conclusion Presence of acetylene as an undesired side product in olefin plants even in small quantities can lead to formation of explosive compounds in the polymerization unit. Hence, it is desired to remove acetylene primarily. Hydrogenation process over

Pd/Ag/α -Al2 O3 catalyst in the tail-end reactor is currently running the art of acetylene removal from ethylene rich streams. Catalyst deactivation occurs due to coke deposition in the hydrogenation catalyst and the regeneration step is inevitable. Hence, any improvement in the catalyst lifetime and performance is greeted by the industries. Accordingly, a new configuration was proposed in the present study. It was managed to use total capacity of each catalyst bed in a configuration consisting 4 beds. A mathematical model based on component and energy balance and catalyst deactivation rate was applied. The accuracy of model was verified against some data from a domestic olefin plant. The conventional and new configurations were compared through acetylene, ethane and ethylene profiles, ethylene yield and total volume of the regenerated catalyst in a 4-year period of

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Fig. 12. Comparison of average overall ethylene yield in runs 1–3.

Fig. 14. Comparing volume of the regenerated catalysts and average overall ethylene yield in 4 years for the conventional and proposed configurations.

operation. Total acetylene elimination was achieved in both configurations. However, a slight increase in ethylene yield was observed in the new configuration. A remarkable decrease in the volume of the regenerated catalyst was observed. The decreased volume of the regenerated catalyst in 4 years was estimated to be 33.3%. Such an achievement could lead to economic benefits for the plants due to reducing cost of regeneration step. Besides, catalyst lifetime and performance are improved in this way. Acknowledgments The authors are grateful to the Shiraz University for supporting this research. Invaluable assistance from the domestic petrochemical plant is also gratefully acknowledged. Appendix A. Heat capacities of the components

Fig. 13. Temperature versus catalyst volume in lead and guard beds in the first and last days of operation in runs 1–3.

Table 11 Situation of each catalyst bed during a 4-year (1440 days) period of operation in the conventional and new configurations.

Heat capacity (J/kmol K)

C p = C1 + C2 [ sinhC3(C/T/T ) ]2 + C4 [ coshC5(C/T /T ) ]2

Component

C1 × 10−5

Acetylene Ethylene Ethane H2 Mixture

0.3199 0.5424 0.3338 0.9479 0.4033 1.3422 0.2762 0.0956  C p = NC i=1 yi C p,i

3

C2 × 10−5

5

C3 × 10−3

C4 × 10−5

C5

1.5940 1.5960 1.6555 2.4660

0.4325 0.5510 0.7322 0.0376

607.10 740.80 752.87 567.60

References

Proposed configuration Run no.

Bed no. 1

Run 1 Lead bed Run 2 Off-line Run 3 Regeneration Run 4 Guard bed Run 5 Lead bed Total operation period Actual configuration Run 1 Lead bed Run 2 Regeneration Run 3 Lead bed Run 4 Regeneration Total operation period

Runtime 2

3

4

(day)

Guard bed Lead bed Regeneration Off-line Guard bed

Off-line Guard bed Lead bed Off-line Regeneration

Off-line Off-line Guard bed Lead bed Regeneration

360 210 314 264 289 4 years

Guard bed Regeneration Guard bed Regeneration

Off-line Lead bed Regeneration Lead bed

Off-line Guard bed Regeneration Guard bed

360 360 360 360 4 years

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