# Application of a tubular flow reactor for oil shale pyro-gasification.

1. Introduction

Energy and chemical application of oil shales using a technology based on pyrolysis and gasification is cost effective and environmentally friendly. Additionally, the chemical and energy potential of the fuel can be utilized to the maximum degree. The current research provides the key concepts of the main pyro-gasification processes:

1) high-speed heating of the processed fuel (103-104 K/s) to control chemical transformations, which involves processing pulverized fuel with a particle size of 0.25-0.35 mm;

2) limiting the maximum temperature to avoid softening of the fuel ash;

3) limiting the residence time for fuel particles in the reaction zone to fractions of a second to avoid unwanted secondary reactions;

4) controlling the process by introducing additives into the fuel dissociation zone, where these are primarily gaseous, oxidizing and reductive;

5) quenching gases recovered through pyro-gasification to maximize the amount of target components.

The above requirements for the pyro-gasification process can be achieved using tubular flow reactors operating on a "gas-suspension" basis. One of the key advantages of tubular reactors is the possibility of supplying large amounts of heat from the external source into the reaction zone through the wall. Figure 1 shows the designed experimental unit [1] for the thermal treatment of pulverized oil shale with a horizontal tubular reactor placed in a fluidized bed of inert materials.

The dried oil shale within the gas flow is transported via reactors where it is heated to the temperature [t.sub.p] needed for the processing. Thus the heat is received through the reactor wall from the fluidized bed. If necessary, the additional heat required for the process can be obtained using the carrier gas to oxidize a certain amount of organic matter from the transported shale. Upon exiting the reactor the gas suspension separates into phases inside the separator. The gas which is the target product from the oil shale processing is transported from the separator and is directed to quenching and further processing. The separated coke is partially burned in the furnace. The greatest amount of coke is utilized for external purposes. The combustion gases coming from the furnace are used as a coolant and fluidizing agent in the high-temperature fluid bed.

[FIGURE 1 OMITTED]

A more detailed description of the principles of operation of the unit is given in [1].

Air (for thermo-oxidative pyrolysis), steam, recycled gas or other gases can be used as the carrier gas in the tubular reactor. The choice of carrier gas depends on the source of oil shale and the desired product. High-speed heating of shale gas suspension in the tubular reactor and rapid cooling of the steam end-products in the quenching heat exchanger, as well as the possibility of regulating the temperature level and residence time of the two streams in the area of heat treatment, can effectively manage the process and thereby influence the composition of the resulting products.

In a complex set of processes that take place in the tubular reactor, heat transfer processes are the primary, and to a great extent influence the operational performance of the entire system. When calculating the heat transfer it is important to take into account the occurrence of chemical reactions in the flow of the fuel gas suspension, and the variability of the flow properties, including the other distinguishing features of thermo-destructive transformations of the original fuel in the flow reactor. Moreover, you need to have relationships which characterize the kinetic and technological parameters of the pyrolysis process.

2. Thermokinetic and technological parameters of pyrolysis

According to experimental data [2], thermo-oxidative pyrolysis of the shale dust occurs in the kinetic response region; therefore, the decomposition rate of the oil shale kerogen under a given level of air consumption is completely determined by the temperature of the process [t.sub.p]. Kashirskij [2] describes oil shales of the Volga basin (dry weight - 58.47% ash, 13.29% carbonate, 28.47% kerogen), which underwent oxidative pyrolysis using an externally heated tube with the diameter D = 16 mm and the length L = 3.8 m. The shale particle size [d.sub.p] was 0-0.250 mm with the initial, as-fired moisture content [W.sup.r] [less than or equal to] 6%, and air was the carrier gas. It was found that for the shale particle sizes [d.sub.p] < 0.300 mm, active kerogen decomposition at a heating rate d[t.sub.p]/d[tau] of about [10.sup.3] [degrees]C/s starts at a temperature of 200-250[degrees]C. The increase in the amount of the gas phase flowing along the reactor during the pyrolysis of oil shale is close to linear. The curve describing the increase in gas consumption is similar to that depicting the temperature increase of the gas suspension flowing along the reactor.

Due to using carbon steel as material of reactor, which can significantly reduce its manufacturing costs, the selection process temperature [t.sub.p] = 600-750 [degrees]C.

Analysis of the gathered data [2] showed the following dependences for the thermo-kinetic and technological pyrolysis characteristics of Kashpirsky oil shale (one of the oil shale fields of the Volga basin) within the temperature range of 600-750[degrees]C:

1) specific yield of pyrolysis gas:

[g.sub.g] = [g.sub.v] x (9.4 x [10.sup.-5] x [t.sub.p] -0.04) x [t.sup.0.613.sub.p] x (1.66 + 0.046/[g.sup.1.77.sub.v]), (1)

where [g.sub.g] is the specific yield of pyrolysis gas, kg/kg of dry oil shale; [g.sub.v] is the specific consumption of feed air, kg/kg of dry oil shale; [t.sub.p] is the process temperature, [degrees]C;

2) specific yield of coke from the reactor:

[g.sub.c] = 0.95 + 0.6 x [g.sub.v] x (1 - [g.sub.g]), (2)

where [g.sub.c] is the specific yield of coke from the reactor, kg/kg of dry oil shale.

The thermal effect of pyrolysis reactions was calculated in accordance with the Hess law as the difference in the amount of heat between the final and initial reaction products, using data from [2] on the component composition of the gas suspension flowing at the inlet and outlet of the reactor. The results of calculating the heat effect in the temperature range 670-680[degrees]C are approximated by:

[q.sub.r] = 1634 x [g.sub.v], (3)

where [q.sub.r] is the thermal effect of pyrolysis reactions, kJ/kg of dry oil shale.

This defines the specific value of the exothermic heat effect at [g.sub.v] [greater than or equal to] 0.04 kg/per kg of dry oil shale. The dependence of [q.sub.r] on [g.sub.v] in Equation (3) is shown in Figure 2. It can be seen that the value of [q.sub.r] is positive, i.e. the reaction is exothermic, and only in case when [V.sub.v] < 30 l/kg of dry oil shale does [q.sub.r] change negative and the reaction becomes endothermic.

[FIGURE 2 OMITTED]

According to [3], the amount of heat released in the combustion reactions of fuel required for its complete combustion air [g.sub.v] is:

[q.sub.com] = 4913 x [g.sub.v], (4)

where [q.sub.com] is the amount of heat released during combustion reactions, kJ/kg of fuel.

It is found that [q.sub.r] < [q.sub.com].

Based on the heat balance of the pyrolysis process the formula for heat to be delivered through the wall of the reactor is as follows:

[??] = 1.65 x [g.sub.v] x [t.sub.p] + 1.3 x [g.sub.c] x [t.sub.p] - 1.1 x [t.sub.T,0] - [g.sub.v] x [t.sub.v,0] - [q.sub.r], (5)

where [??] is the amount of heat which must go through the wall of the reactor, kJ/kg of dry oil shale; [t.sub.T,0] and [t.sub.v,0] are the initial temperatures of shale and air, respectively, [degrees]C.

3. External and internal heat transfer

The heat transfer coefficient of the fluidized bed reactor with horizontal pipes, for fluidization velocity values [w.sub.wor] [greater than or equal to] [w.sub.opt], where [w.sub.opt] is the optimum speed (corresponding to the maximum rate of heat transfer), can be calculated according to the Zabrodsky formula [4], as shown by Equation (6):

[[alpha].sub.k.si.] = 35.8 x [[lambda].sup.0.6.sub.p.c.] x [[rho].sup.0.2.sub.p] x [d.sup.-0.36.sub.p,k.sl.], (6)

where [[alpha].sub.k.sh.] is the heat transfer coefficient of the fluidized bed reactor with horizontal pipes, W/([m.sup.2] x K); [[lambda].sub.p.c] is the thermal conductivity of coke combustion products, W/(m x K); [[rho].sub.p] is the density of solids, kg/[m.sup.3]; [d.sub.p,k.sl.] is the size of particles in a fluidized bed, m.

If substituting the numerical values of the fluidizing gas properties and particles, the result is as follows:

[[alpha].sub.k.sl.] = 43 x [d.sup.-0.36.sub.p,k.sl.]. (7)

When placing the tube bundle with a relative pitch of S/D [greater than or equal to] 2.5 within 0.2 m above the gas distribution grid to H (height of the original dense layer), heat transfer patterns are the same as for single tubes [4]. A relatively dense arrangement of pipes is possible without reducing the heat transfer.

The specific area of the fluidized bed section per kg/s of dry processed shale is:

[f.sub.k.sl.] = [F.sub.k.sl.]/[G.sub.s] = [g.sub.p.c.]/[[[rho].sub.p.c.] x [w.sub.wor]]

= 5.18 x [10.sup.-9] x [??] x [(1050 - [t.sub.k.sl.]).sup.-1] x (18 + 5.5 x [10.sup.6] x [d.sup.1.5.sub.p,k.sl.]) x [d.sup.-2.sub.p, k.sl.], (8)

where [f.sub.k.sl.] is the specific cross-sectional area of the fluidized bed, [m.sup.2] x s/kg of dry oil shale; [F.sub.k.sl.] is the surface area of the fluidized bed, [m.sup.2]; [G.sub.s] is the mass flow of the solid phase, kg/s; [g.sub.p.c.] is the specific consumption of coke combustion, kg/kg of dry oil shale; [[rho].sub.p.c.] is the density of the products of combustion of coke, kg/[m.sup.3]; [w.sub.wor] is the working fluidization velocity, m/s; [t.sub.k.sl.] is the temperature of the fluidized bed, [degrees]C.

The specific surface area of heat transfer reactor tubes in a fluidized bed, per 1 kg/s of dry processed shale:

f = F/[G.sub.s] = [[10.sup.3] x [??]]/[q.sub.w] = [10.sup.3] x [??]/[[[alpha].sub.k.sl.] x ([t.sub.k.sl.] - [t.sub.w])], (9)

where f is the heat transfer surface area of the reactor tubes in a fluidized bed, [m.sup.2] x s/kg of dry oil shale; F is the integral (total) heat exchange surface of the reactor tubes, [m.sup.2]; [t.sub.w] is the pipe wall temperature, [degrees]C; [q.sub.w] is the heat flux at the wall, W/[m.sup.2].

Using (7), Equation (9) will acquire the form:

f = 23.26 x [??] x [d.sup.0.36.sub.p,k.sl.] x [([t.sub.k.sl] - [t.sub.w]).sup.-1]. (10)

Dividing (10) by (8) term by term, we have:

F/[F.sub.k.sl.] = f/[f.sub.k.sl.] = 4.485 x [10.sup.9] x [d.sup.2.36.sub.p.k.sl.] x (1050 - [t.sub.k.sl.])

x [[([t.sub.k.sl.] - [t.sub.w]) x (18 + 5.5 x [10.sup.6] x [d.sup.1.5.sub.p,k.sl.])].sup.-1]. (11)

The ratio F /[F.sub.k.sl.] can be determined by the structural characteristics. The height of the particle bed prior to its fluidization should be no less than 0.6 m to ensure the necessary degree of fluidized bed uniformity. The heating surface is at the height corresponding to the dense layer. In this case maximum heat transfer coefficients are provided which are the same for all rows of posted reactor tubes.

Calculation of the internal heat transfer in the reactor tube is performed using similarity equations for the average heat flow of gas suspension to the decomposition of solid particles [5], and a numerical method to estimate the heat transfer of the chemically reacting gas suspension flow [6].

4. Calculation results

An algorithm has been developed for various pyrolysis calculations. Some of the results of calculations related to reactor diameters of 0.02 and 0.04 m are shown in Table 1.

The data in Table 1 show that feeding the reactor with wet slate requires an increase in the length of the reactor. The amount of coke, [g.sub.c], produced from oil shale pyrolysis under the conditions given in Table 1 equals 0.75-0.85 kg/per kg of dry oil shale. The need for coke combustion in the furnace process is significantly lower at 0.184 kg/per kg of dry oil shale at [t.sub.p] = 600[degrees]C and 0.208 kg/per kg of dry oil shale at 700[degrees]C. Thus the process of oxidative pyrolysis of oil shale described in [1] is enclosed in terms of heat energy.

Heat exchange between the shale gas suspension flow and the wall is of high intensity. For example, at the initial concentration of particles [K.sub.0] = 5.1 (kg/h)/(kg/h), the air velocity [w.sub.0] = 11 m/s, the reactor diameter D = 0.02 m, and the average particle diameter [d.sub.p] = 0.150 mm, the average heat transfer coefficient, [alpha], of the reacting flow of gas suspension equals 180 W/([m.sup.2] x K). Under the same conditions, the average heat transfer coefficient, a, of the flow of clean air without particles equals 83 W/([m.sup.2] x K). The curves depicting the change in the values of parameters along the reactor under the above process conditions are shown in Figure 3.

The data in Figure 3 are typical and show that the temperature of the reactor wall increases with the length of the stream reactor, x. The difference between temperatures of the wall at the output and input of the reactor under the considered conditions may amount to 100[degrees]C or more.

Almost over the entire length of the reactor during the heating the flow temperature [t.sub.p,aw], is below the temperature [t.sub.g,aw]. As seen from Figure 3, at the end of the reactor the temperature [t.sub.p,aw]. becomes higher than [t.sub.g,aw]. The temperature rises with an increase in [V.sub.v]. This is the result of exothermic oxidation reactions of shale particles during the heating process.

A significant decrease in the concentration of particles K along the length of the stream reactor x (Fig. 3) is due to gasification. The sliding speed factor of the particles, [[phi].sub.v], increases at the initial part of the reactor where rapid warming of the flow occurs, and slowly decreases along the length of the reactor, x.

In Figure 3 parameters (dashed lines) for the chemically inert flow at [Q.sub.r] = 0 and those of the reacting flow are provided, so they can be compared.

Data on external heat exchange are shown in Figure 4, which depicts the change of the working fluidization velocity and temperature of the fluidized bed (Fig. 4a), as well as parameters F/[F.sub.k.sl.] and [f.sub.k.sl.] (Fig. 4b) with the particle diameter [d.sub.p,k.sl.] On the computational model adopted the change of the particle diameter [d.sub.p,k.sl.] does not affect the length of the reactor and its heat output. This effect is mediated via the temperature of the fluidized bed, [t.sub.k.sl.], which increases with [d.sub.p,k.sl.] (Fig. 4a). The correlation of F/[F.sub.k.sl.] and [f.sub.k.sl.], with the particle diameter [d.sub.p,k.sl.] (Fig. 4b) allows management of the design dimensions and layout of the fluidized bed reactor in the layer by varying [d.sub.p,k.sl.].

[FIGURE 3 OMITTED]

[FIGURE 4 OMITTED]

For the particle size [d.sub.p,k.sl.] = 6 mm, as an example, the following space arrangement can be considered. With a fluidized bed height of 1.8 m and an area of the gas distribution grid of 5.3 x 1 m calculations yielded [S.sub.1] = 3D and [S.sub.2] = 0.866[S.sub.1] for the reactor diameter D = 0.02 m and number of reactors N = 630, and the surface area of heat transfer [F.sub.[SIGMA]] = 212 [m.sup.2].

At the oil shale flow rate [G.sub.s,0] = 0.02 kg/s in one reactor, the capacity of the unit based on processed fuel is 46 t/h or 1104 t/day. At [t.sub.p] = 600[degrees]C the amount of pyrolysis gas produced [G.sub.g] is 400 t/day ([V.sub.v] = 100 l/kg of dry oil shale) and coke (net supply to the furnace) [G.sub.c] is 618 t/day. The thermal output from the fluidized bed will be close to 5 MW and the total heat output, including the heat of pyrolysis reactions, is 8 MW.

Table 2 compares the characteristics of contemporary processes of gasification of solid fuels [7-9] and thermooxidative Volga shale pyrolysis in tubular reactors. It can be seen that the consumption of processed oil shale per unit reactor volume and pyrolysis gas yield per unit cross-sectional area of the reactor for tubular reactors is one-two orders of magnitude higher than that for the other known reactor devices with dense or fluid layers.

5. Conclusions

The residence time for the oil shale particles in the reaction zone of tubular reactors is within the range 0.3-0.4 s, which distinguishes the gas-suspended pyrolysis mode from long-residence time processes with a dense layer of particles or fluid. In contrast to those devices, high-speed pyrolysis in tubular reactors is manageable. The tubular reactors provide the needed heat input via the tube walls from an external source, and they produce a quality gas product which is not diluted with other components.

The specific parameters, such as consumption of processed oil shale per reactor volume and yield of pyrolysis gas per cross section of reactor, are much higher for tubular reactors than for existing reactor devices with dense and fluid layers.

SYMBOLS: D - tube diameter, m; dp - size of solid particles, equal to the diameter of the ball, which is equivalent to the particle surface, m; F - surface area, [m.sup.2]; G - mass flow, kg/s; g - specific consumption, kg/kg of dry oil shale; [V.sub.v] - air flow, l/kg of dry oil shale; K--expenditure mass concentration of shale particles in the gas stream, (kg/h)/(kg/h); Q - capacity of the heat flow, W; q - heat flux, W/[m.sup.2]; t - temperature, [degrees]C; w - average speed of the continuous phase in the section of the channel, m/s; x - length of the reactor, m; [alpha] - heat transfer coefficient, W/([m.sup.2] x K); [lambda] - thermal conductivity, W/(m x K); [rho] - density, kg/[m.sup.3]; [tau] - time, s; [[phi].sub.v] = u/w - particle velocity slip factor; u - average speed solids section of the channel, m/s.

INDICES: v - air; g - gaseous phase; s - solid phase; w - parameter on the wall or at a wall temperature; av. - option when the average temperature of; r - chemical reaction; p - flow, process; p.c. - gaseous products of combustion of coke; k.sl. - fluidized bed.

doi: 10.3176/oil.2014.3.04

REFERENCES

[1.] Kosova, O. Yu. Installation for thermal pulverized oil shale treatment. Oil shale as an alternative source for fuel and raw materials. Fundamental research. Experience and Prospects. Proceedings of International Scientific Conference, Saratov, May 21-23, 2007, Saratov State Technical University, Saratov, 2007, 108-112 (in Russian).

[2.] Kashirskij, V. G. Experimental Basics of a Complex Energy and Technological Usage of Fuels. Saratov State University Publishers, Saratov, 1981, 144 pp (in Russian).

[3.] Baskakov, A. P., Berg, B. V., Witt, O. C. et al. Heat Engineering. Energoatomizdat, Moscow, 1991, 224 pp (in Russian).

[4.] Tishchenko, A. T., Khvastukhin, Yu. I. Furnaces and Fluidized Bed Heat Exchangers. Naukova Dumka, Kiev, 1973, 146 pp (in Russian).

[5.] Pechenegov, Yu. Y. Heat transfer and hydraulic resistance under gas suspension flow accompanied by solid phase gasification. II Russian National Conference on Heat Transfer, Vol. 5. Two-phase types of flow. Dispersed types of flow and porous media, Moscow, October 26-30, 1998. Publishing House Moscow Power Engineering Institute, Moscow, 1998, 260-262 (in Russian).

[6.] Pechenegov, Yu. Y., Kosova, O. Yu. Method for calculating the heat flow in a gas suspension pipe with a thermochemically decomposed solid phase. Problems of gas dynamics and heat transfer in power plants. Proceedings of XIV Summer School for Young Scientists Supervised by Academician A. I. Leontyeva, Rybinsk, Yaroslavl region, May 26-30, 2003. Vol. 1. Publishing House Moscow Power Engineering Institute, Moscow, 2003, 306-308 (in Russian).

[7.] Volkov, E. P., Gavrilov, N. F. A promising technology for the use of low-grade fuels. Izv. RAN. Energetika, 2005, 3, 135-147 (in Russian).

[8.] Pechuro, N. S., Kapkin, V. D., Pessin, O. Yu. Chemistry and Technology of Synthetic Liquid Fuels and Gas. Khimiy, Moscow, 1986, 352 pp (in Russian).

[9.] Chemical Technology of Solid Fuels (Makarov, G. N., Kharlampovich, G. D., eds). Khimiy, Moscow, 1986, 496 pp (in Russian).

Presented by A. Siirde

Received October 17, 2013

YURY Y. PECHENEGOV, VENJAMIN F. SIMONOV, BORIS A. SEMYONOV, OLGA YU. KOSOVA, ANTON N. MRAKIN *

Faculty of Power Engineering, Yuri Gagarin State Technical University of Saratov, 77 Politechnitcheskaya street, Saratov, Russia, 410054

* Corresponding author: e-mail anton1987.87@mail.ru

Energy and chemical application of oil shales using a technology based on pyrolysis and gasification is cost effective and environmentally friendly. Additionally, the chemical and energy potential of the fuel can be utilized to the maximum degree. The current research provides the key concepts of the main pyro-gasification processes:

1) high-speed heating of the processed fuel (103-104 K/s) to control chemical transformations, which involves processing pulverized fuel with a particle size of 0.25-0.35 mm;

2) limiting the maximum temperature to avoid softening of the fuel ash;

3) limiting the residence time for fuel particles in the reaction zone to fractions of a second to avoid unwanted secondary reactions;

4) controlling the process by introducing additives into the fuel dissociation zone, where these are primarily gaseous, oxidizing and reductive;

5) quenching gases recovered through pyro-gasification to maximize the amount of target components.

The above requirements for the pyro-gasification process can be achieved using tubular flow reactors operating on a "gas-suspension" basis. One of the key advantages of tubular reactors is the possibility of supplying large amounts of heat from the external source into the reaction zone through the wall. Figure 1 shows the designed experimental unit [1] for the thermal treatment of pulverized oil shale with a horizontal tubular reactor placed in a fluidized bed of inert materials.

The dried oil shale within the gas flow is transported via reactors where it is heated to the temperature [t.sub.p] needed for the processing. Thus the heat is received through the reactor wall from the fluidized bed. If necessary, the additional heat required for the process can be obtained using the carrier gas to oxidize a certain amount of organic matter from the transported shale. Upon exiting the reactor the gas suspension separates into phases inside the separator. The gas which is the target product from the oil shale processing is transported from the separator and is directed to quenching and further processing. The separated coke is partially burned in the furnace. The greatest amount of coke is utilized for external purposes. The combustion gases coming from the furnace are used as a coolant and fluidizing agent in the high-temperature fluid bed.

[FIGURE 1 OMITTED]

A more detailed description of the principles of operation of the unit is given in [1].

Air (for thermo-oxidative pyrolysis), steam, recycled gas or other gases can be used as the carrier gas in the tubular reactor. The choice of carrier gas depends on the source of oil shale and the desired product. High-speed heating of shale gas suspension in the tubular reactor and rapid cooling of the steam end-products in the quenching heat exchanger, as well as the possibility of regulating the temperature level and residence time of the two streams in the area of heat treatment, can effectively manage the process and thereby influence the composition of the resulting products.

In a complex set of processes that take place in the tubular reactor, heat transfer processes are the primary, and to a great extent influence the operational performance of the entire system. When calculating the heat transfer it is important to take into account the occurrence of chemical reactions in the flow of the fuel gas suspension, and the variability of the flow properties, including the other distinguishing features of thermo-destructive transformations of the original fuel in the flow reactor. Moreover, you need to have relationships which characterize the kinetic and technological parameters of the pyrolysis process.

2. Thermokinetic and technological parameters of pyrolysis

According to experimental data [2], thermo-oxidative pyrolysis of the shale dust occurs in the kinetic response region; therefore, the decomposition rate of the oil shale kerogen under a given level of air consumption is completely determined by the temperature of the process [t.sub.p]. Kashirskij [2] describes oil shales of the Volga basin (dry weight - 58.47% ash, 13.29% carbonate, 28.47% kerogen), which underwent oxidative pyrolysis using an externally heated tube with the diameter D = 16 mm and the length L = 3.8 m. The shale particle size [d.sub.p] was 0-0.250 mm with the initial, as-fired moisture content [W.sup.r] [less than or equal to] 6%, and air was the carrier gas. It was found that for the shale particle sizes [d.sub.p] < 0.300 mm, active kerogen decomposition at a heating rate d[t.sub.p]/d[tau] of about [10.sup.3] [degrees]C/s starts at a temperature of 200-250[degrees]C. The increase in the amount of the gas phase flowing along the reactor during the pyrolysis of oil shale is close to linear. The curve describing the increase in gas consumption is similar to that depicting the temperature increase of the gas suspension flowing along the reactor.

Due to using carbon steel as material of reactor, which can significantly reduce its manufacturing costs, the selection process temperature [t.sub.p] = 600-750 [degrees]C.

Analysis of the gathered data [2] showed the following dependences for the thermo-kinetic and technological pyrolysis characteristics of Kashpirsky oil shale (one of the oil shale fields of the Volga basin) within the temperature range of 600-750[degrees]C:

1) specific yield of pyrolysis gas:

[g.sub.g] = [g.sub.v] x (9.4 x [10.sup.-5] x [t.sub.p] -0.04) x [t.sup.0.613.sub.p] x (1.66 + 0.046/[g.sup.1.77.sub.v]), (1)

where [g.sub.g] is the specific yield of pyrolysis gas, kg/kg of dry oil shale; [g.sub.v] is the specific consumption of feed air, kg/kg of dry oil shale; [t.sub.p] is the process temperature, [degrees]C;

2) specific yield of coke from the reactor:

[g.sub.c] = 0.95 + 0.6 x [g.sub.v] x (1 - [g.sub.g]), (2)

where [g.sub.c] is the specific yield of coke from the reactor, kg/kg of dry oil shale.

The thermal effect of pyrolysis reactions was calculated in accordance with the Hess law as the difference in the amount of heat between the final and initial reaction products, using data from [2] on the component composition of the gas suspension flowing at the inlet and outlet of the reactor. The results of calculating the heat effect in the temperature range 670-680[degrees]C are approximated by:

[q.sub.r] = 1634 x [g.sub.v], (3)

where [q.sub.r] is the thermal effect of pyrolysis reactions, kJ/kg of dry oil shale.

This defines the specific value of the exothermic heat effect at [g.sub.v] [greater than or equal to] 0.04 kg/per kg of dry oil shale. The dependence of [q.sub.r] on [g.sub.v] in Equation (3) is shown in Figure 2. It can be seen that the value of [q.sub.r] is positive, i.e. the reaction is exothermic, and only in case when [V.sub.v] < 30 l/kg of dry oil shale does [q.sub.r] change negative and the reaction becomes endothermic.

[FIGURE 2 OMITTED]

According to [3], the amount of heat released in the combustion reactions of fuel required for its complete combustion air [g.sub.v] is:

[q.sub.com] = 4913 x [g.sub.v], (4)

where [q.sub.com] is the amount of heat released during combustion reactions, kJ/kg of fuel.

It is found that [q.sub.r] < [q.sub.com].

Based on the heat balance of the pyrolysis process the formula for heat to be delivered through the wall of the reactor is as follows:

[??] = 1.65 x [g.sub.v] x [t.sub.p] + 1.3 x [g.sub.c] x [t.sub.p] - 1.1 x [t.sub.T,0] - [g.sub.v] x [t.sub.v,0] - [q.sub.r], (5)

where [??] is the amount of heat which must go through the wall of the reactor, kJ/kg of dry oil shale; [t.sub.T,0] and [t.sub.v,0] are the initial temperatures of shale and air, respectively, [degrees]C.

3. External and internal heat transfer

The heat transfer coefficient of the fluidized bed reactor with horizontal pipes, for fluidization velocity values [w.sub.wor] [greater than or equal to] [w.sub.opt], where [w.sub.opt] is the optimum speed (corresponding to the maximum rate of heat transfer), can be calculated according to the Zabrodsky formula [4], as shown by Equation (6):

[[alpha].sub.k.si.] = 35.8 x [[lambda].sup.0.6.sub.p.c.] x [[rho].sup.0.2.sub.p] x [d.sup.-0.36.sub.p,k.sl.], (6)

where [[alpha].sub.k.sh.] is the heat transfer coefficient of the fluidized bed reactor with horizontal pipes, W/([m.sup.2] x K); [[lambda].sub.p.c] is the thermal conductivity of coke combustion products, W/(m x K); [[rho].sub.p] is the density of solids, kg/[m.sup.3]; [d.sub.p,k.sl.] is the size of particles in a fluidized bed, m.

If substituting the numerical values of the fluidizing gas properties and particles, the result is as follows:

[[alpha].sub.k.sl.] = 43 x [d.sup.-0.36.sub.p,k.sl.]. (7)

When placing the tube bundle with a relative pitch of S/D [greater than or equal to] 2.5 within 0.2 m above the gas distribution grid to H (height of the original dense layer), heat transfer patterns are the same as for single tubes [4]. A relatively dense arrangement of pipes is possible without reducing the heat transfer.

The specific area of the fluidized bed section per kg/s of dry processed shale is:

[f.sub.k.sl.] = [F.sub.k.sl.]/[G.sub.s] = [g.sub.p.c.]/[[[rho].sub.p.c.] x [w.sub.wor]]

= 5.18 x [10.sup.-9] x [??] x [(1050 - [t.sub.k.sl.]).sup.-1] x (18 + 5.5 x [10.sup.6] x [d.sup.1.5.sub.p,k.sl.]) x [d.sup.-2.sub.p, k.sl.], (8)

where [f.sub.k.sl.] is the specific cross-sectional area of the fluidized bed, [m.sup.2] x s/kg of dry oil shale; [F.sub.k.sl.] is the surface area of the fluidized bed, [m.sup.2]; [G.sub.s] is the mass flow of the solid phase, kg/s; [g.sub.p.c.] is the specific consumption of coke combustion, kg/kg of dry oil shale; [[rho].sub.p.c.] is the density of the products of combustion of coke, kg/[m.sup.3]; [w.sub.wor] is the working fluidization velocity, m/s; [t.sub.k.sl.] is the temperature of the fluidized bed, [degrees]C.

The specific surface area of heat transfer reactor tubes in a fluidized bed, per 1 kg/s of dry processed shale:

f = F/[G.sub.s] = [[10.sup.3] x [??]]/[q.sub.w] = [10.sup.3] x [??]/[[[alpha].sub.k.sl.] x ([t.sub.k.sl.] - [t.sub.w])], (9)

where f is the heat transfer surface area of the reactor tubes in a fluidized bed, [m.sup.2] x s/kg of dry oil shale; F is the integral (total) heat exchange surface of the reactor tubes, [m.sup.2]; [t.sub.w] is the pipe wall temperature, [degrees]C; [q.sub.w] is the heat flux at the wall, W/[m.sup.2].

Using (7), Equation (9) will acquire the form:

f = 23.26 x [??] x [d.sup.0.36.sub.p,k.sl.] x [([t.sub.k.sl] - [t.sub.w]).sup.-1]. (10)

Dividing (10) by (8) term by term, we have:

F/[F.sub.k.sl.] = f/[f.sub.k.sl.] = 4.485 x [10.sup.9] x [d.sup.2.36.sub.p.k.sl.] x (1050 - [t.sub.k.sl.])

x [[([t.sub.k.sl.] - [t.sub.w]) x (18 + 5.5 x [10.sup.6] x [d.sup.1.5.sub.p,k.sl.])].sup.-1]. (11)

The ratio F /[F.sub.k.sl.] can be determined by the structural characteristics. The height of the particle bed prior to its fluidization should be no less than 0.6 m to ensure the necessary degree of fluidized bed uniformity. The heating surface is at the height corresponding to the dense layer. In this case maximum heat transfer coefficients are provided which are the same for all rows of posted reactor tubes.

Calculation of the internal heat transfer in the reactor tube is performed using similarity equations for the average heat flow of gas suspension to the decomposition of solid particles [5], and a numerical method to estimate the heat transfer of the chemically reacting gas suspension flow [6].

4. Calculation results

An algorithm has been developed for various pyrolysis calculations. Some of the results of calculations related to reactor diameters of 0.02 and 0.04 m are shown in Table 1.

The data in Table 1 show that feeding the reactor with wet slate requires an increase in the length of the reactor. The amount of coke, [g.sub.c], produced from oil shale pyrolysis under the conditions given in Table 1 equals 0.75-0.85 kg/per kg of dry oil shale. The need for coke combustion in the furnace process is significantly lower at 0.184 kg/per kg of dry oil shale at [t.sub.p] = 600[degrees]C and 0.208 kg/per kg of dry oil shale at 700[degrees]C. Thus the process of oxidative pyrolysis of oil shale described in [1] is enclosed in terms of heat energy.

Heat exchange between the shale gas suspension flow and the wall is of high intensity. For example, at the initial concentration of particles [K.sub.0] = 5.1 (kg/h)/(kg/h), the air velocity [w.sub.0] = 11 m/s, the reactor diameter D = 0.02 m, and the average particle diameter [d.sub.p] = 0.150 mm, the average heat transfer coefficient, [alpha], of the reacting flow of gas suspension equals 180 W/([m.sup.2] x K). Under the same conditions, the average heat transfer coefficient, a, of the flow of clean air without particles equals 83 W/([m.sup.2] x K). The curves depicting the change in the values of parameters along the reactor under the above process conditions are shown in Figure 3.

The data in Figure 3 are typical and show that the temperature of the reactor wall increases with the length of the stream reactor, x. The difference between temperatures of the wall at the output and input of the reactor under the considered conditions may amount to 100[degrees]C or more.

Almost over the entire length of the reactor during the heating the flow temperature [t.sub.p,aw], is below the temperature [t.sub.g,aw]. As seen from Figure 3, at the end of the reactor the temperature [t.sub.p,aw]. becomes higher than [t.sub.g,aw]. The temperature rises with an increase in [V.sub.v]. This is the result of exothermic oxidation reactions of shale particles during the heating process.

A significant decrease in the concentration of particles K along the length of the stream reactor x (Fig. 3) is due to gasification. The sliding speed factor of the particles, [[phi].sub.v], increases at the initial part of the reactor where rapid warming of the flow occurs, and slowly decreases along the length of the reactor, x.

In Figure 3 parameters (dashed lines) for the chemically inert flow at [Q.sub.r] = 0 and those of the reacting flow are provided, so they can be compared.

Data on external heat exchange are shown in Figure 4, which depicts the change of the working fluidization velocity and temperature of the fluidized bed (Fig. 4a), as well as parameters F/[F.sub.k.sl.] and [f.sub.k.sl.] (Fig. 4b) with the particle diameter [d.sub.p,k.sl.] On the computational model adopted the change of the particle diameter [d.sub.p,k.sl.] does not affect the length of the reactor and its heat output. This effect is mediated via the temperature of the fluidized bed, [t.sub.k.sl.], which increases with [d.sub.p,k.sl.] (Fig. 4a). The correlation of F/[F.sub.k.sl.] and [f.sub.k.sl.], with the particle diameter [d.sub.p,k.sl.] (Fig. 4b) allows management of the design dimensions and layout of the fluidized bed reactor in the layer by varying [d.sub.p,k.sl.].

[FIGURE 3 OMITTED]

[FIGURE 4 OMITTED]

For the particle size [d.sub.p,k.sl.] = 6 mm, as an example, the following space arrangement can be considered. With a fluidized bed height of 1.8 m and an area of the gas distribution grid of 5.3 x 1 m calculations yielded [S.sub.1] = 3D and [S.sub.2] = 0.866[S.sub.1] for the reactor diameter D = 0.02 m and number of reactors N = 630, and the surface area of heat transfer [F.sub.[SIGMA]] = 212 [m.sup.2].

At the oil shale flow rate [G.sub.s,0] = 0.02 kg/s in one reactor, the capacity of the unit based on processed fuel is 46 t/h or 1104 t/day. At [t.sub.p] = 600[degrees]C the amount of pyrolysis gas produced [G.sub.g] is 400 t/day ([V.sub.v] = 100 l/kg of dry oil shale) and coke (net supply to the furnace) [G.sub.c] is 618 t/day. The thermal output from the fluidized bed will be close to 5 MW and the total heat output, including the heat of pyrolysis reactions, is 8 MW.

Table 2 compares the characteristics of contemporary processes of gasification of solid fuels [7-9] and thermooxidative Volga shale pyrolysis in tubular reactors. It can be seen that the consumption of processed oil shale per unit reactor volume and pyrolysis gas yield per unit cross-sectional area of the reactor for tubular reactors is one-two orders of magnitude higher than that for the other known reactor devices with dense or fluid layers.

5. Conclusions

The residence time for the oil shale particles in the reaction zone of tubular reactors is within the range 0.3-0.4 s, which distinguishes the gas-suspended pyrolysis mode from long-residence time processes with a dense layer of particles or fluid. In contrast to those devices, high-speed pyrolysis in tubular reactors is manageable. The tubular reactors provide the needed heat input via the tube walls from an external source, and they produce a quality gas product which is not diluted with other components.

The specific parameters, such as consumption of processed oil shale per reactor volume and yield of pyrolysis gas per cross section of reactor, are much higher for tubular reactors than for existing reactor devices with dense and fluid layers.

SYMBOLS: D - tube diameter, m; dp - size of solid particles, equal to the diameter of the ball, which is equivalent to the particle surface, m; F - surface area, [m.sup.2]; G - mass flow, kg/s; g - specific consumption, kg/kg of dry oil shale; [V.sub.v] - air flow, l/kg of dry oil shale; K--expenditure mass concentration of shale particles in the gas stream, (kg/h)/(kg/h); Q - capacity of the heat flow, W; q - heat flux, W/[m.sup.2]; t - temperature, [degrees]C; w - average speed of the continuous phase in the section of the channel, m/s; x - length of the reactor, m; [alpha] - heat transfer coefficient, W/([m.sup.2] x K); [lambda] - thermal conductivity, W/(m x K); [rho] - density, kg/[m.sup.3]; [tau] - time, s; [[phi].sub.v] = u/w - particle velocity slip factor; u - average speed solids section of the channel, m/s.

INDICES: v - air; g - gaseous phase; s - solid phase; w - parameter on the wall or at a wall temperature; av. - option when the average temperature of; r - chemical reaction; p - flow, process; p.c. - gaseous products of combustion of coke; k.sl. - fluidized bed.

doi: 10.3176/oil.2014.3.04

REFERENCES

[1.] Kosova, O. Yu. Installation for thermal pulverized oil shale treatment. Oil shale as an alternative source for fuel and raw materials. Fundamental research. Experience and Prospects. Proceedings of International Scientific Conference, Saratov, May 21-23, 2007, Saratov State Technical University, Saratov, 2007, 108-112 (in Russian).

[2.] Kashirskij, V. G. Experimental Basics of a Complex Energy and Technological Usage of Fuels. Saratov State University Publishers, Saratov, 1981, 144 pp (in Russian).

[3.] Baskakov, A. P., Berg, B. V., Witt, O. C. et al. Heat Engineering. Energoatomizdat, Moscow, 1991, 224 pp (in Russian).

[4.] Tishchenko, A. T., Khvastukhin, Yu. I. Furnaces and Fluidized Bed Heat Exchangers. Naukova Dumka, Kiev, 1973, 146 pp (in Russian).

[5.] Pechenegov, Yu. Y. Heat transfer and hydraulic resistance under gas suspension flow accompanied by solid phase gasification. II Russian National Conference on Heat Transfer, Vol. 5. Two-phase types of flow. Dispersed types of flow and porous media, Moscow, October 26-30, 1998. Publishing House Moscow Power Engineering Institute, Moscow, 1998, 260-262 (in Russian).

[6.] Pechenegov, Yu. Y., Kosova, O. Yu. Method for calculating the heat flow in a gas suspension pipe with a thermochemically decomposed solid phase. Problems of gas dynamics and heat transfer in power plants. Proceedings of XIV Summer School for Young Scientists Supervised by Academician A. I. Leontyeva, Rybinsk, Yaroslavl region, May 26-30, 2003. Vol. 1. Publishing House Moscow Power Engineering Institute, Moscow, 2003, 306-308 (in Russian).

[7.] Volkov, E. P., Gavrilov, N. F. A promising technology for the use of low-grade fuels. Izv. RAN. Energetika, 2005, 3, 135-147 (in Russian).

[8.] Pechuro, N. S., Kapkin, V. D., Pessin, O. Yu. Chemistry and Technology of Synthetic Liquid Fuels and Gas. Khimiy, Moscow, 1986, 352 pp (in Russian).

[9.] Chemical Technology of Solid Fuels (Makarov, G. N., Kharlampovich, G. D., eds). Khimiy, Moscow, 1986, 496 pp (in Russian).

Presented by A. Siirde

Received October 17, 2013

YURY Y. PECHENEGOV, VENJAMIN F. SIMONOV, BORIS A. SEMYONOV, OLGA YU. KOSOVA, ANTON N. MRAKIN *

Faculty of Power Engineering, Yuri Gagarin State Technical University of Saratov, 77 Politechnitcheskaya street, Saratov, Russia, 410054

* Corresponding author: e-mail anton1987.87@mail.ru

Table 1. Reactor parameters and process conditions of the oxidative pyrolysis of Volga basin oil shales D, m L, m [g.sub.v], kg/kg [W.sup.r], [t.sub.p], [t.sub.0], of dry oil shale % [degrees]C [degrees]C 0.02 12.00 0.065 0 600 214 0.02 5.30 0.130 0 600 232 0.02 4.70 0.130 0 600 274 0.02 3.00 0.195 0 600 192 0.02 3.20 0.195 0 700 367 0.04 9.40 0.162 0 600 242 0.04 7.70 0.195 0 600 192 0.04 6.10 0.195 0 600 253 0.04 8.24 0.195 6 600 171 0.04 9.30 0.195 12 600 128 D, m [t.sub.k.sl], [G.sub.S,0], [K.sub.0], [g.sub.g], kg/kg [degrees]C kg/s kg/kg of dry oil shale 0.02 722 0.047 15.4 0.40 0.02 692 0.020 7.7 0.36 0.02 683 0.018 7.7 0.36 0.02 704 0.013 5.1 0.40 0.02 773 0.009 5.1 0.50 0.04 682 0.070 6.4 0.38 0.04 693 0.061 5.1 0.40 0.04 679 0.053 5.1 0.40 0.04 698 0.064 5.1 0.40 0.04 709 0.071 5.1 0.40 D, m [w.sub.0], Q, kW [Q.sub.r], m/s kW 0.02 13.4 23.72 5.62 0.02 11.8 7.92 4.66 0.02 11.7 6.33 4.3 0.02 11.0 5.1 4.65 0.02 10.8 3.75 3.34 0.04 12.9 26.1 19.91 0.04 12.4 23.7 21.1 0.04 12.4 18.6 16.2 0.04 12.5 26.0 22.1 0.04 12.5 32.9 24.6 Indicated: D and L--diameter and length of the reactor; [W.sup.r]--moisture of the source shale; [t.sub.0] - temperature of the flow at the intake; [t.sub.ksl]--temperature of the fluidized bed; [G.sub.s,0] - consumption of oil shale; Q--total heat of the process; Qr--reaction heat. Table 2. Key parameters for the current gasification of solid fuels and thermooxidative pyrolysis of the Volga basin oil shale in tubular reactors Process Data Single [d.sub.p], [t.sup.p], productivity for mm [degrees]C installation on fuel, t/h Lurgi (dense layer of coal 5-30 1100 steam-oxygen blowing, 40-75 P = 20-25 am) Winkler (fluidized coal [less than or 1000 bed, steam-oxygen or 20-35 equal to] 10 air blast, P = 1 am) Koppers-T ottsek coal [less than or 1500 (dust-coal flame, to 40 equal to] 0.1 steam-oxygen blast, P = 1 am) UTT-3000 oil shale [less than or 500 (P = 1 am) 139 equal to] 20 Tubular reactors oil shale (air blast, 46 0.15 600 P = 1 am): D = 0.02 m; N= 630 pcs; D = 0.04 m; N = 315 pcs 69.2 0.15 600 Process Data Residence Gas Heat at gas time for output, combustion, particles in thousand kJ/[m.sup.3] the reactor, s [m.sup.3]/h Lurgi (dense layer of 5000 75 12000-16300 steam-oxygen blowing, P = 20-25 am) Winkler (fluidized 100-500 60 7500-9400 bed, steam-oxygen or air blast, P = 1 am) Koppers-T ottsek ca 1 50 10300-11700 (dust-coal flame, steam-oxygen blast, P = 1 am) UTT-3000 ca 2000 5-8 48400 (P = 1 am) Tubular reactors (air blast, 0.3 16.7 12000 P = 1 am): D = 0.02 m; N= 630 pcs; D = 0.04 m; N = 315 pcs 0.4 27.7 11000 Process Data Thermal Specific fuel stress consumption, reaction volume, kg/([m.sup.3]-h) GJ/([m.sup.3]-h) Lurgi (dense layer of 2.9-5.4 250 steam-oxygen blowing, P = 20-25 am) Winkler (fluidized 0.9-8.4 71 bed, steam-oxygen or air blast, P = 1 am) Koppers-T ottsek 3.6-25 360 (dust-coal flame, steam-oxygen blast, P = 1 am) UTT-3000 - 463 (P = 1 am) Tubular reactors (air blast, 27.2 4822 P = 1 am): D = 0.02 m; N= 630 pcs; D = 0.04 m; N = 315 pcs 18 2563 Process Data Intensity of Volume of the gas, the reactor kg/([m.sup.2]-h) apparatus, [m.sup.3] Lurgi (dense layer of 2000 160 steam-oxygen blowing, P = 20-25 am) Winkler (fluidized 3000 550 bed, steam-oxygen or air blast, P = 1 am) Koppers-T ottsek 3000 60 (with 4 (dust-coal flame, nozzles) steam-oxygen blast, P = 1 am) UTT-3000 400 300 (P = 1 am) Tubular reactors (air blast, 230000 9.5 P = 1 am): D = 0.02 m; N= 630 pcs; D = 0.04 m; N = 315 pcs 180121 27

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Author: | Pechenegov, Yury Y.; Simonov, Venjamin F.; Semyonov, Boris A.; Kosova, Olga Yu.; Mrakin, Anton N. |
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Publication: | Oil Shale |

Geographic Code: | 4EXRU |

Date: | Sep 1, 2014 |

Words: | 4522 |

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