To help introduce the basics of SPC, we'll assume that the variable being monitored is the specific gravity (SG) of n-hexane as it is being produced. We'll assume that the SG is measured four times per day at 0300, 0900, 1500, and 2100 by the plant's laboratory. Table 1 shows the results for a three day time period. A-grade industrial n-hexane must have a SG between 0.61 and 0.69. First, we'll see how the process engineer analyzes this data.

SPC Overview

In a continuous chemical process, two types of charts are commonly used: individual value or X-bar charts and moving range (MR) or R-bar charts. X-bar charts are used on a regular basis to monitor the process during a time of change. For example, an R-bar chart would be appropriate if you were changing the feeds to the process. The R-bar chart weights more recent data more heavily than historical data.

The chemical industry typically uses one of two types of process control. 3-sigma control specifies quality limits nearly equal to process limits. 6-sigma control specifies quality limits that are twice as large as control limits. We'll focus on the 3-sigma system.

With all of the different types of limits, it's easy to become confused. For our n-hexane process, we'll have 6 different limits we'll consider. Three UCL's (Upper Control Limits) and three LCL's (Lower Control Limits).

UCL (calculated) = statistical upper control limit
UCL (process) = pre-determined, acceptable process upper control limit
UCL (quality) = pre-determined, acceptable quality upper control limit

LCL (calculated) = statistical lower control limit
LCL (process) = pre-determined, acceptable process lower control limit
LCL (quality) = pre-determined, acceptable quality lower control limit

The process limits are those which define boundaries of operation for the process or an acceptable operating value. The quality control limits are those used to "grade" material. The term "quality limits" will refer to the A grade or top grade material limits. You should realize that there are also B and C grades of materials that companies often sell as well. The limits of these other grades vary accordingly. Essentially, the farther away from specifications a product is, the lower the grade, and its value decreases sharply.

Typically in a 3-sigma system, the process limits are said to be "tighter" than the quality limits by 5-10%. This is done so that even if the process exceeds process limits by a small amount, it will still be within quality standards. However, the 6-sigma system dictates that the process limits be half of the quality limits. For example, if you had an upper quality control limit of 100, the upper process control limit in a 6-sigma system would be 50 while a 3-sigma system may have an upper process control limit of around 90. Basically, a 6-sigma system requires more strict (and sometimes unrealistic) control, depending on the process. This is why many chemical manufactures implement the 3-sigma system. Now that we've discussed the different types of limits and charts involved, let's see how our system is performing!

SPC: X-bar Charts

Eq. (1)

Eq. (2) Eq. (3) Eq. (4) Eq. (5)

In a 3-sigma system, Z is equal to 3 [Equations 2 and 3] (hence its name) and a 6-sigma system uses Z=6, therefore:
UCL (calculated) = 0.65 + 3(0.0056) = 0.67
LCL (calculated) = 0.65 - 3(0.0056) = 0.63

As mentioned before the other control limits are set depending on the quality of the product needed. A-grade n-hexane must be between 0.61 and 0.69 (these are the quality limits). Typical process limits may then be 0.62 to 0.68. Now we know all of the limits for our current data:
UCL (calculated) = 0.67
UCL (process) = 0.68
UCL (quality) = 0.69

LCL (calculated) = 0.63
LCL (process) = 0.62
LCL (quality) = 0.61

At this point, it's tempting to conclude that since the calculated limits are inside the process and quality limits, the process is operating perfectly. But let's have a look at the X-bar chart:

Figure 1 shows the performance from 2/25/99 to 2/28/99. According to the definition of "in control", the process should meet four criteria:
1. No sample points outside of process limits
2. Most points near average
3. Nearly equal number of points above and below average
4. Points are randomly distributed
According to Figure1, only conditions 3 and 4 are being met. This process should be examined for process upsets or interruptions in stability. After the appropriate process changes were made, another X-bar chart was constructed over another 4 day period, Figure 2 below shows these results:

After the process improvements, the data suggests that the process is in control and all four criteria for control are being met. Figure 2 shows how you should aim to control your process.

SPC: R-bar Charts

R-bar charts utilize chart factors that are typically found in statistical references. Table 2 shows a portion of such a chart for 3-sigma control:

The UCL (calculated) are LCL (calculated) are defined by:

Eq. (6) Eq. (7) Eq. (8)

where MR (moving range) is the absolute value of the difference between the current data point and the preceding data point. The number of sub-groups is an area that most people do not agree upon. For example, if you group results by week and you're analyzing data for a month you could use 4 sub-groups. As a general rule, if your continuous process has been operating under the same specifications over the time of your analysis, you may assume 2 sub-groups. This is the approach we'll use for our system.

Let's assume that our plant also produces glycol which has an average specific gravity of 1.11. An R-bar chart provides an effective means of monitoring the transition from n-hexane (SG=0.65) to glycol (SG=1.11). Monitoring the individual results (X-bar chart) in conjunction with the R-bar chart will paint a very clear picture of the transition. The data in Table 3 shows the 4-day transition. The feeds were changed just prior to this data being recorded. What we're seeing is the n-hexane leaving the system and the glycol showing up gradually.

Now, to form the R-bar chart, we graph the data/time versus the moving range. We can assume 2 sub-groups (n-hexane and glycol over a short period of time). The moving range for 3/15/99 at 0900 was calculated by |(0.63-0.65)|=0.02. Typically, there will be no process or quality control limits for R-bar charts. For this transition:

Figure 3 shows the R-bar chart for the transition from n-hexane to glycol. While using an X-bar chart alone during times of change in a system is feasible, it is sometimes difficult to graph the data due to the potentially large differences in results. Another consideration is that in data compiling, a R-bar chart appearing between X-bar charts is a nice way to show a transition phase has occurred. R-bar charts are also useful when plotting data over a large time span. To show the contrast in the two types of charts, the x-bar chart during the transition is shown in Figure 4.

The arrow in Figure 4 shows the point at which the SG of the glycol has entered the quality control range. This is a very important point because the system output must be directed to a glycol storage tank at this point (rather than the waste container used during the transition). This is why the charts should be used in conjunction with one another, and an R-bar chart should not be used alone during a transition.

Uses of R-bar charts include:
1. Keep a record of when process changes or feed changes occurred
2. Record of how long the process took to stabilize
3. Show long history of a process or piece of equipment

If or whenever you use R-bar charts, remember that they tell you nothing about the actual value of the results, only deviations from one result to the next.

SPC: Danger Signs and Where to Start Looking

As a process or quality engineer, you'll eventually come across some charts that make your eyes pop out and spell "O-V-E-R-T-I-M-E". Imagine going to work one morning and find the chart below:

Depending on where the results came from, the problem could be several things.
If the results are from online measuring devices:
1. Check the operator's log for any abnormal behavior during the time that the results were out of standard.
2. Check the calibration schedule for the measuring device.

If the results are from a laboratory:
1. Check laboratory equipment for correct calibration.
2. Review laboratory notes on the tests for any errors that may have been made in the testing procedure.
3. Ask the laboratory technician if he/she remembers anything strange about the tests. For example, sample collection container abnormalities that may have led to contamination.

Any investigating beyond these ideas may begin to be counterproductive. However, the process should be monitored closely to see if this result is repeated later. The process has returned to normal operation and the out of standard results were not a serious compromise of quality.

Now suppose you find a chart resembling Figure 6:

You may initially think that since no results are out of the quality control standards that you don't have a problem. To the contrary, unless there has been an intentional process change, you have a very serious problem. It's not a matter of where the process has been or where it is now, but where it is going. This type of trend cannot be attributed to simple error, there is something seriously wrong! Depending on many factors, you must find a place to start investigating. You may want to start with equipment that can immediately affect the SG. Look closely at the process data over the past few days. It can be helpful to compare data over a comparable period of time when the process was in control versus this new trend that you're seeing. I might suggest starting with the separation equipment.

Until this point we've considered only one characteristic of n-hexane, specific gravity. Process engineers must simultaneously monitor all important characteristics of a product. Suppose your manager presents you with the following two charts:

You notice that the SG seems to be fine while the concentration has dropped off dramatically. While SG is generally a good indicator of concentration, it doesn't appear to be so in this situation. Since you have no reason to doubt the accuracy of these results at this time, where should you start? The first question you should answer is: "What else is in the stream that is lowering the concentration?" A quick look at the gas chromatograph show that on 3/4/99 at 1500 (when the n-hexane concentration was at 92%) there was also significant amounts of two other chemicals: 2-methyl-butene-3 (SG=0.63) and isoprene (SG=0.68)[hypothetical components]. Since the 2-methyl-butene-3 is lighter than n-hexane and isoprene is heavier than n-hexane, the contaminant mixture did not force the specific gravity out of standard, but the concentration is being seriously affected (this is why you monitor both SG and concentration). Now all you have to do is find out how it got there. I might start with heat exchangers that may be leaking.

Ranging from an early level of 20 Gcal/tonne (79.4 MBtu/tonne) to about 7 Gcal/tonne (27.8 MBtu/tonne) in the last decade of the 20th century. {parse block="google_articles"}The energy intensive nature of the process is the key driving force for improving the technology and reducing the overall cost of manufacturing Looking further ahead, we'll review some potentially significant developments and concepts that may impact the manner in which ammonia is produced. Some of these manufacturing routes are being tested or employed at a few plants around the world, but have yet to be fully developed into commercial processes. We'll also review more traditional approaches to ammonia manufacturing along the way.

Reforming Section Developments

In the conventional process, steam reforming is carried out in a fired furnace of the side fired or top fired type. Both need large surface areas for uniform heat distribution along the length of the catalyst tubes.

Figure 1: Overall Layout of a Steam Reforming Plant for Ammonia Synthesis (Click to Enlarge)

This process has several disadvantages. For example, it is a thermally inefficient process (about 90% including the convection zone) and there are mechanical and maintenance issues. The process is difficult to control and reforming plants require a large capital investment.

Gas Heated Reformers

Future technologies include the use of Gas Heated Reformers (GHR), which are tubular gas-gas exchangers. In the GHR, the secondary reformer outlet gases supply the reforming heat. Though it is not presently being used widely, GHR has certain advantages over fired furnaces. Table 1 shows a list of these advantages. Kellogg's Reforming Exchanger System is an example of GHR technology. Although GHR results in reduced energy consumption, a comprehensive energy conservation network should be established to maximize the benefits of a GHR system.

Lechatelier's Principle states that a reaction equilibrium can be shifted by applying external forces. This offers a means of removing products from the reaction mixture to increase the conversion per pass. In reforming, experiments have been performed up to 500 °C (932 ^°F) and 20 bar (294 psig) using a palladium membrane to remove the product hydrogen. These experiments have results in a significant increase in methane conversion as can be seen by the following case study.

The reduced air requirement (about 80% of conventional plants) in the secondary reformer is required with a 60% hydrogen removal rate in the reformer. This will also require an additional source of nitrogen. Therefore, the technologies in which nitrogen is being added separately, either from an Air Separation Unit (ASU) or from any other sources, will become more important in this case.

Isobaric Manufacturing

The primary hurdle in the isobaric method of manufacturing ammonia is the poor conversion of methane at elevated pressure. The bottleneck is the maximum permissible temperature range due to metallurgical constraints in the reformer tubes. Synthesis pressures are no longer an issue with the development of the Kellogg Advanced Ammonia Process (KAAP), which utilizes a ruthenium-based catalyst operating at 90-100 ata (1470 psia). Thus, if the methane conversion can be increased by hydrogen separation, the process can be operated at higher isobaric pressures.

The synthesis compressor can be reduced to one small compressor at the natural gas feed. The power consumption in this compressor will be 3.0 MW for an isobaric pressure of 100 ata in the front end because of reduced gas flow and reduced differential pressure. The gas flow in synthesis compressor remains near 200,000 Nm³/h (117,715 scfm) while the flow will be reduced to near 45,000 Nm³/h (26,485 scfm) in natural gas compressor. The differential pressure in the synthesis compressor is 175 kg/cm²a (from 25 kg/cm²g to 200 kg/cm²g), while the differential pressure is only 60 kg/cm² a in natural gas compressor (from 40 kg/cm²g to 100 kg/cm²g). The power consumption is around 3.0 MW in the conventional recirculator. This will result in a total power consumption of 6.0 MW in raising the pressure of synthesis gas. Presently, the power consumption in the synthesis gas compressor is around 25.0 MW for a conventional ammonia plant at same load. This ,however, requires the process air compressor to be operated at a discharge pressure of 100 ata (1470 psia) compared with a pressure of 34-35 ata (510 psia) in the conventional plant. The net energy saving in the isobaric process can be near 0.5 Gcal/tonne (1.98 MBtu/tonne). Moreover, it will also save the energy in CO₂ compressor of the urea plant because the CO₂ from the ammonia plant will be available at a much higher pressure.

Shift Section Developments

The water-gas shift reaction is favorable for producing carbon dioxide which is used as a raw material for urea production. Presently, most plants use a combination of conventional High/Low Temperature Shift (HTS/LTS) or High/Medium/Low Temperature Shift (HTS/MTS/LTS) technology. Another option is a combination of HTS/LTS/Selectoxo technology. While not as common as the other combinations, this arrangment offers advantages that will be discussed later. The most important objectives for this section are a low pressure drop and efficient heat recovery from the process gas.{parse block="google_articles"}

Selectoxo Unit

The Selectoxo unit offers several advantages for grass root designs as well as for revamps. Selectoxo (or selective catalytic oxidation) was developed by Engelhard for oxidizing carbon monoxide while not oxidizing hydrogen. The Selectoxo process provides good energy efficiency because it minimizes carbon moxide "slip" (only about 0.03%), improved process flexibility, and higher productivity in revamps when compared to other oxidation options. The Selectoxo unit is capable of increasing a plant's capacity by 1.5-2.0%.

The Selectoxo unit can also help in a grass root plant by maintaining carbon dioxide/ammonia production ratios which is favorable for full conversion of ammonia to urea. The economics of this option are to be considered against the extra cost of carbon dioxide production by other means (either from the flue gas of the primary reformer or through back burning of extra synthesis gas).

Carbon Dioxide Removal Section Developments

The removal of carbon dioxide has been performed via solvent absorption and distillation since the inception of ammonia technology processes. This section of the ammonia plant is the largest consumer of energy after the cooling water system. The energy consumption is due to thermally inefficient distillation, dissipation of huge amounts of low level heat into the cooling water via product carbon dioxide, and pressurization and depressurization of absorbents.{parse block="google_articles"}

Isobaric Manufacturing

Chemical absorption in the isobaric manufacturing of ammonia can be unattractive because of the very high pressure (100 ata). Therefore, major changes in the existing carbon dioxide removal technologies may be necessary. Replacement technologies may include cryogenic condensation or pressure swing absorption (PSA).

Carbon dioxide separation through PSA is offered in the Low Cost Ammonia Process (LCA). PSA is scalable an may be more economical because of efficient carbon dioxide recovery at higher pressures. However, further development in this direction is essential for the recovery of high purity carbon dioxide as desired in urea production.

Carbon dioxide separation via condensation may also become more attractive due to an increased concentration of carbon dioxide which can be realized with successful hydrogen separation through membranes. This would allow the concentration of carbon dioxide to be increased by 18 to 36 mole percent. This would allow carbon dioxide concentrations in the gas to be reduced to 15% by chilling of the 100 ata fron end gases. This method also provides high pressure carbon dioxide for urea production which will reduce the power consumption in the carbon dioxide compressor of the urea plant substantially. The remaining product carbon dioxide gas can be recovered via PSA. A combined PSA and condensation process may solve the problem of carbon dioxide purity from the PSA process.

Final Purification of Synthesis Gas

The conventional methanation process can result in the loss of hydrogen. Minimizing this loss is of prime concern when examining the process used to purify the syngas.

Pressure Swing Absorption (PSA) Unit

PSA represents an effective means of reducing the hydrogen loss in the methanator. In this process, the product hydrogen is separated out from the raw synthesis gas and then nitrogen is added. The other benefit is the production of pure synthesis gas, which saves on recycle compression and the elimination of the losses through the purge gas stream by way of eliminating the purge itself.{parse block="google_articles"}

Cryogenic Separation Process

Cryogenic separation of inert gases from the raw synthesis gas is a commonly used approach. This unit is integrated into the purge gas recovery loop from the back to the front end of the ammonia unit. It serves to recover hydrogen from the purge stream and feed it back to the ammonia synthesis loop after recompression.

In this separation process, inerts in the synthesis gas are removed through cryogenic condensation. Typically, the composition of conventionally prepared synthesis gas is about 74% hydrogen, 0.8-1.0% methane, 0.32% argon with the balance being nitrogen. In this process, nearly all of the methane is removed along with half of the argon present, thus it produces "cleaner" synthesis gas for ammonia production. Moreover, the hydrogen to nitrogen ratio of the synthesis gas can be controlled independently without affecting the performance of front end. Traditionally, this ratio is controlled by varying the process air flow to the secondary reformer which makes the system reactive between front end and the back end. A cryogenic separation unit eliminates the dependence of the back end on the performance of the front end.

However, this process does not contribute to energy savings. Rather, it represents a good option for revamps after achieving the limits of capacity using conventional revamps. The cryogenic separation process creates additional margin in the front end by allowing more methane slip and by reducing the total quantity of inerts in the loop.

Ammonia Synthesis

Several developments in ammonia synthesis have been made in the past, these developments revolve around the basic principles of reactioin, heat recovery, cooling, production ammonia separation, and recycling of synthesis gas.

Synthesis Catalysts

After almost 90 years of a monopoly in the ammonia synthesis market, iron catalyst has not been replaced by a precious metal (ruthenium) based catalyst used in the KAAP developed by Kellogg. The KAAP catalyst is reported to be 40% more active than iron catalysts.{parse block="google_articles"}

Research work on low temperature and low pressure catalysts to produce ammonia at 20-40 kg/cm²g and 100 ^°C is being performed at Project and Development India Ltd. (PDIL) according to their in-house magazine. The catalyst being studied is based on cobalt and ruthenium metals and has exhibited few encouraging results.

Ammonia Separation

The removal of product ammonia is accomplished via mechanical refrigeration or absorption/distillation. The choice is made by examining the fixed and operating costs. Typically, refrigeration is more economical at synthesis pressures of 100 ata or greater. At lower pressures, absorption/distillation is usually favored. A comparison of these two methods is presented in Table 3.

Minimizing the amount of ammonia in the recycle gas of an ammonia process presents an interesting scenario. Usually the ammonia concentration of the recycle is 3-4%, but reducing this amount to 1.5% can increase plant capacity by about 2.5%. However, the additional separation can often represent a significant addition to the capital cost of the plant and may not be economical for retrofitting (depending on operating pressure). However, reduced ammonia concentration in the recycle can be reviewed for a grass root project where capacity gains can be realized with an additional investment.

Decreasing the ammonia concentration in the recycle stream of existing plants is usually hampered by the high energy cost required for water absorption. Norsk Hydro (Norway) developed a method of reducing the recycle ammonia concentration to near 0.5% via absorption in glycol (DEG). This process can be installed in a high pressure loop (>100 ata) and in combination with a condensation unit. The installed cost is said to be lower than a comparable mechanical refrigeration system.

The separation of product ammonia within the converter using liquid or solid adsorbent can increase the system efficiency significantly. The regenerated adsorbent is fed to the converter and contacts the reaction mixture. Product ammonia is absorbed and removed from the converter. The product ammonia can be recovered either by changing the pressure or temperature depending on process economics. This method would eliminate the need for a synthesis loop and the recycling of synthesis gas. This concept is still being investigated in academic research.

Final Word

The developments discussed here such as isobaric manufacturing, the use of gas heat reformers, hydrogen separation, carbon dioxide removal technology, product ammonia separation, and high activity synthesis catalyst can result in a significant reduction in energy consumption when compared with traditional technology.

Global demand, increased competition, and ingenuity have fueled efforts to enhance existing ammonia technology. In an industry where change is often accepted reluctantly, these technological advancements will have to prove themselves worthy before receiving industry-wide attention.

At this time, the new standards for hydrochloric acid scrubber efficiency are not known, but we can be sure that they will be more rigorous than present requirements. At present, most fume scrubbers are designed for efficiencies of 90-95% - new requirements are expected to be in the 98-99% range. What does this mean for the pickle line operator?{parse block="google_articles"}

Why Scrubbers are Needed

The liquid hydrochloric acid used in pickling is actually a solution of hydrogen chloride gas in water. Raw acid is generally purchased as ‘20°Baumé’, which means it has a specific gravity of 1.16 and contains about 32% hydrogen chloride (HCl) and 68% water. For pickling purposes, this acid is usually diluted to about 18-22% HCl.

The HCl solution used in pickling emits HCl vapors into the space above the liquid surface. The tendency of the HCl to escape is measured as a vapor pressure - the higher this pressure, the more HCl escapes. Figure 1 shows the vapor pressure curves for HCl at various temperatures as a function of liquid concentration - the vapor pressure increases as temperatures and concentrations increase. These curves are for the vapor pressure over solutions in pure water - the vapor pressures over pickling acid solutions are significantly higher, due to the accumulation of iron chloride in the acid, which increases the solution HCl vapor pressure.

Figure 1: Vapor Pressure of Hydrogen Chloride Over Aqueous Solutions (ref 1,2)

Figure 1 shows why pickle acids need fume control. The threshold limit value (TLV) for HCl is 5 ppmv, which is equivalent to a vapor pressure of about 0.004 mmHg. This pressure is in equilibrium with a 5% solution at 120°F, 2% at 140°F and less than 1% at 176°F - all concentrations well below what is used in pickling, whether in open or closed tanks. This means that the air above the surface of the acid during pickling is well above the allowable limits, and some form of containment or control is needed.

Besides operator hygiene, control is needed because HCl is extremely corrosive to the building fabric and equipment. Referring again to Figure 1, the 80°F curve shows that there is negligible vapor pressure up to about 5% - this means that HCl is very soluble in cold water, and even low gas concentrations can generate high liquid acid concentrations. Thus, HCl in the air can migrate through the interior fabric of the building, and dissolve in condensation on cold surfaces, producing a strong and corrosive acid in hidden or inaccessible areas.

The ideal fume control would be to pickle in a sealed vessel, isolated from the air - however, the need for openings to insert, guide, access and remove the steel from the acid generally precludes this, so usually the fumes are controlled by creating a draft above the acid surface. This causes air to be drawn in through the openings, thereby preventing escape of acid fumes. However, as the air flows across the acid surface, it sweeps away the HCl vapors, which are then carried out of the tank in the exhaust air. The concentration of HCl in the exhaust gas is a complex function of tank open surface area, air velocity, temperature and acid composition, and can vary from 50 ppmv in well-controlled open tank systems with lateral exhaust, to as much as 2000 ppmv in continuous pickle lines with tight lids and 6000 ppmv in tower picklers. These concentrations are far too high for discharge to the atmosphere, even without considering the effect of such emissions on the building roof and nearby structures.

Scrubber Design

{parse block="google_articles"}There are as many types of scrubber as there are inventors - but they all make use of the physical chemistry shown in Figure 1. By contacting the HCl bearing gas with cold water, the fumes are dissolved, and removed from the air. This is illustrated in Figures 2 and 3. In Figure 2, at point ‘A’, which represents conditions in the pickle tank, the HCl vapor pressure over 10% acid at 180°F is 0.86 mmHg, whereas the air entering the tank has no HCl in it (if the fume control system is working properly). Thus, the tendency is for the HCl to evaporate into the air. In Figure 3, at point ‘B’, which represents conditions in the scrubber, the air from the pickle tank, now containing perhaps 500 ppmv of HCl (equal to 0.4 mmHg pressure), is contacted with water that has a negligible HCl vapor pressure, and the tendency is for the HCl to dissolve in the water.

Figure 2: Conditions in the Pickle Tank

Figure 3: Conditions in the Fume Scrubber

When HCl dissolves in water, a substantial amount of heat is generated - in a scrubber, this heat is absorbed by evaporation of water into the air. The evaporation of water will also cool the incoming air to near the adiabatic saturation temperature, so that the air leaving a wet scrubber will have almost 100% humidity.

Although there are numerous scrubber designs, three types dominate the market - packed, plate and crossflow.

The packed scrubber is by far the most widely used type. Figure 4 shows the construction of a packed scrubber. The air enters the bottom of the scrubber, and flows upwards through a bed of packing. Scrubbing water is sprayed on top of the packing bed, and flows downwards by gravity - the purpose of the packing is to promote turbulence, and good mixing of the liquid and the gas, so that the HCl will be absorbed.

Figure 4 shows the components of a packed scrubber , which are the shell, the packing support, the packing, the liquid distribution system and the demister. The shell is usually plastic - PVC, polypropylene or FRP - and is most economically constructed as a cylinder. At the bottom is a grid which supports the packing, and which needs to have a large open area to allow passage of the air and gas in opposite directions, without creating excessive pressure drop.

The packing consists of plastic shapes, typically 2" to 4" in size. There are numerous proprietary packing designs - some common ones are: rings, which are short hollow cylinders; saddle shaped packing; donuts of small dia. rod wound in a spiral; and intersecting plates, usually with a spherical external profile. Each packing claims superiority, backed by lab and pilot testing - in fact, they all have about the same efficiency under industrial conditions.

Figure 4: Components of a Packed Fume Scrubber

Figure 5: Components of a Plate Fume Scrubber

Figure 6: Components of a Cross-Flow Fume Scrubber

Good liquid distribution on the top of the packing is essential - if more liquid is added in one area of the packing, it runs down that area, while the gas, taking the path of least resistance travels up areas where there is less liquid, and contacting is poor.

Finally, the droplets entrained in the air stream must be removed, in order to stop them being discharged from the stack as acid rain. This is done by passage through a device which changes flow direction, and provides a surface for the removed droplets to grow, in order to prevent re-entrainment. The two main types of entrainment eliminator are: chevron, consisting of parallel S-shaped blades; and mesh, which is a bed of knitted polypropylene fibers. They are both equally efficient, but the chevron is preferred, because it lasts longer, has large openings, a lower pressure drop, and less tendency to block with dust, crystals or deteriorated fibers.

The advantages of packed scrubbers are simplicity and cheapness (mainly due to traditional use of cheap materials and flimsy construction methods), plus the ability to accommodate variations in air flow. Disadvantages are the tendency of the packing to plug with dirt and the need for a large water flow over the packing (about 100 gal water/10,000 cu.ft. gas). Usually, to avoid high water consumption, and generation of large amounts of very weak acid effluent, this water flow is obtained by pumping water from the bottom of the scrubber back to the top of the packing. This requires a pump to be maintained, and also puts acidic water, instead of clean water, in contact with the clean air.

The components of a plate scrubber are shown in Figure 5. The shell and demister are the same as for a packed scrubber, but the gas/liquid contact is brought about by a number of perforated plates. The plates are plastic plates, perforated with numerous small holes, 1/8" to 3/8" in diameter. As the air passes through these holes at high velocity, it prevents the water from falling through the holes, and generates a highly turbulent water/air mixture on the plate. The water depth on the plate is controlled by weirs - as more water is added to the top plate of the scrubber, it displaces water over the weir, and this flows by gravity through a downcomer pipe to the plate below, eventually discharging from the base of the scrubber.

Plate scrubbers use relatively low amounts of water (1-2 gal water/10,000 cu.ft. gas) on a once-through basis, and require minimal maintenance. The low volume of relatively concentrated effluent makes for easier recovery. However, these scrubbers are not suitable for systems in which the air flow varies widely.

The features of cross-flow scrubbers are shown in Figure 6. These are packed scrubbers in which the gas flow is horizontal instead of vertical - the water still flows downwards by gravity. In these scrubbers, the shell is frequently rectangular in cross section, and the packing is held between two grids. The water from the sump is pumped back to the top of the packing.

These scrubbers have the same advantages and disadvantages as packed scrubbers, but are less efficient. Their main advantage is a low profile when headroom is limited, and the ability to locate the fan on ground level with minimum ducting.

Scrubber Performance and Efficiency

This discussion relates to scrubbers which are operating with effluent acid concentrations of 5% or less, so that the HCl vapor pressure of the liquid can be considered negligible. Tower picklers, where recovered acid up to 15% can be generated, required more complex calculations.{parse block="google_articles"}

The efficiency of removal of HCl in a packed scrubber depends on the air and liquid flow mass velocities, and the size of packing. Generally, the volume of packing needed for a given efficiency is approximately constant, so the same results can be obtained by a short, large diameter, scrubber, or a taller, small diameter scrubber. Large diameter requires a lot more water volume to irrigate the packing - small diameter gives high pressure drop, and the risk of flooding (when the water is held up in the packing due to the high gas velocity). A superficial air velocity of 5-8 fps is usually found to be the best compromise.

In packed scrubbers, the height of the packed bed is given by the equation

Z = H_og x N_og

where H_og is the height of a transfer unit (which is an experimentally determined property of the packing used), based on gas concentrations, and N_og is the number of transfer units, based on gas concentrations. The number of transfer units required, at low HCl concentrations, is:

N_og = ln [1/(1 - E_o)]

where E_o is the desired overall scrubbing efficiency (as a fraction).

The number of transfer units required for various overall efficiencies is given in Table 1.

This shows that, in order to increase the efficiency of a given scrubber from 90 to 99%, the depth of the packing has to be doubled (from 2.3 to 4.6 transfer units), and to get to 99.5%, has to increase 2.3 times (5.3/2.3 = 2.3). This increases the pressure drop through the scrubber proportionally, but does not increase the overall height in the same proportion - usually, the packing accounts for only about 1/2 to 1/3 of the overall scrubber height.

For plate scrubbers, the efficiency of a single plate varies in the range 60-80%, depending on air and water flow velocities, and plate spacing. Usually, it is desirable to minimize the diameter of plate scrubbers, without causing flooding or excessive entrainment, and they are typically designed with air velocities in the range 6-9 fps. The number of plates required for a given overall efficiency is:

n = ln(1 - E_o)/ln(1 - e)

where E_o is the overall efficiency, and e is the efficiency of a single plate, both expressed as fractions.

Table 2 gives the overall efficiency for various numbers of plates at three different plate efficiencies.

Thus, present day plate scrubbers usually have 2 or 3 plates - to meet future efficiency requirements, 4 or 5 plates will be needed. For each additional plate, the height of the scrubber increases by 24-30", and the pressure drop by 1" to 1.5" of water.

Scrubber Operations

The scrubber is only one component of a fume exhaust system, although there is a tendency to look upon it as being the only thing that matters. However, the function of the scrubber is to remove HCl from the air - collection of the fumes is determined by hood and duct design, and the amount of draft is determined by the fan performance.{parse block="google_articles"}

Some areas where the scrubber design is important are:

• variable loads
• entrainment
• stack droplet emissions

The HCl concentration in the scrubber inlet gas can vary substantially, particularly in batch pickling operations, in which high HCl loads can occur as the steel is being removed from the tank - this is especially the case in modern systems that have crane-mounted hoods to control these fumes. Essentially, a fume scrubber is a constant-efficiency device, so, if the inlet concentration goes up 10 times, the outlet concentration also goes up 10 times. Usually, these high concentrations only occur for short periods, so that the average emissions remain low. However, such excursions increase the acidity of any droplets in the gas after the scrubber, and make entrainment control even more important.

In order to prevent acid rain from the stack, water droplets must be eliminated from the air stream. In the scrubber, this is done by having an efficient entrainment eliminator. However, all wet scrubbers discharge air saturated with water vapor, so that contact with cold equipment or stack walls can cause condensation after the scrubber - this condensation absorbs the residual HCl content of the gas, and becomes highly acidic. If the stack velocity is too high, or if the stack wall is very cold, or the scrubber exhaust air temperature high, the condensate may be discharged as droplets from the stack. In order to prevent this, an additional entrainment eliminator may be necessary at the top of the stack. This presents problems in compliance testing, because the rules require testing in the stack, at a point which does not represent the actual stack emission!

Another factor that needs consideration is what to do with the scrubber effluent. The scrubber only removes the HCl from the air, generating a dilute acid effluent stream in the process. This effluent requires treatment or re-use. With plate scrubbers, which use little water, it is possible to generate effluent acid up to 2-3% strength, which can often be returned to the pickle tanks to make up for evaporation - this is not generally possible with packed scrubbers, that generate effluent in the 0.1 to 0.5% concentration range. Recycle of the scrubber effluent makes economic sense - a continuous strip pickle line can discharge acid costing as much as $25,000/yr, and then spend as much again in neutralizing chemicals.

Controlling Costs

Costs for scrubbers are always of concern, and will become even more important as more efficient scrubbers are required. The initial cost of the scrubber can be minimized by designing the fume exhaust system to handle the smallest volume of air that will give satisfactory control of fumes. It is tempting to oversize exhaust systems, just to be safe, and to avoid having to pay too much attention to maintenance. However, oversizing results in higher capital costs, more fan hp, more effluent generation, and higher acid losses from the pickle tanks.

The ideal fume control system is a closed cover, with no exhaust, as in ‘fumeless picklers’ used for wire strands, but this is not technically practical for batch pickling or high speed strip lines. Techniques for minimizing exhaust rates include:

• tight, well-maintained covers
• locate exhaust ducts near openings in hoods and tanks
• minimize open area with local seals or closures
• use of double covers
• regular maintenance of fan, hoods and ducting.
• careful, well-balanced duct design

Conclusions

Future emission regulations are going to require larger, more efficient HCl fume scrubbers. The new limits are easily achievable using scrubbers that are 30-50% larger than those presently in use. The cost of upgrading fume scrubbers can be reduced by improving fume control practices to minimize exhaust rates.

References

1. Perry, R.H., and Green, D.W., ‘Perry’s Chemical Engineers’ Handbook’, 6th ed., McGraw-Hill, New York, 1984, p 3-64
2. Fritz, J.J., and Fuget, C.R., ‘Vapor pressure of aqueous hydrogen chloride solutions, 0 to 50°C’, Ind.Eng Chem., Chem. and Eng. Data, Vol1, #1, p 10, 1956

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Brief History - Enhancing the Methanol-I Plant Capacity

Originally, this plant was designed to operate on feed gas from an ammonia plant consisting of a gas mixture of 75% hydrogen, 22% carbon dioxide, 1% carbon monoxide and some inerts.  The reaction of the methanol in gas rich in CO₂ is milder as it produces water along with methanol.  The crude methanol concentration is also lower. Water further retards the rate of reaction. The two reactions involved here are:

H₂  +  CO₂  ---->    CH₃OH  +  H₂O          +            9.8 kcal/kgmol            

H₂  +  CO    ---->   CH₃OH                     +          21.6 kcal/kgmol                       

Normally, a gas mixture of H₂ + CO + CO₂ is used in a proportion measured in terms of a “R” value (H₂-CO₂)/(CO+CO₂) equal to 2.0 to get an optimum methanol conversion per pass.

Changes to the process included:

1. A methanol chiller was introduced in the gas cooling circuit at the reactor outlet to reduce the methanol concentration and temperature in the recycle gas which helped to increase the methanol production from 4~5 MTD up to 75 MTD level.
2. Setting up a synthesis gas generation unit (SGGU) to supply CO rich gas from natural gas reformer in February 1998.  This gas composition is better for methanol production compared to the rectisol wash gas which is rich in CO₂. The synthesis gas and distillation loops were debottlenecked by replacing of some control valves, installation of exchangers, and other modifications.  The capacity was boosted to 100 to 120 MTD.
3. Replacement of the refining column trays with high capacity Superfrac™ trays from Kostch Glistch India Ltd in October 2002.  This, along with other peripheral modifications, were made to increase distillation capacity to 145 MTD.
4. Replacement of the quench adiabatic methanol convertor to Linde’s Isothermal Reactor and debottlenecking of the distillation loops for higher capacity.  The capacity of the plant was increased to 160 MTD in September 2003.   

Major advantages of Isothermal Reactor include: 

Lower pressure drop in reactor
Less temperature variation
Increased life of catalyst
Narrow band of temperature differences in the reactor catalyst bed
Sustained Production Level throughout the catalyst life due to better conversion
Less by-product formation
Effective heat recovery

Figure 1 below shows the temperature profile in the isothermal reactor.  

Compared to the expected 160 MTD production capacity, the unit has achieved a stable production level of 185~190 MTD.

A flow diagram of the new loop is shown in Figure 2 below.  In this article, we’ll focus on this latest dimension added to the plant, highlighting the re-commissioning experiences. 

	Â		Â
Figure 1: Comparing Reactor Temperature Profiles Before and After the Changes	Â	Figure 2: Changes in the Methanol Synthesis and Distillation Loops	Â	Figure 3: Methanol Synthesis and Distillation Loops After Changes

Plant Re-Commissioning with Isothermal Reactor

Following the replacement of the quench reactor with the Isothermal reactor from Linde, the plant was ready for start up.  The following details the activities associated with start up after the changes were made.{parse block="google_articles"}

Basic and Detail Engineering - Design Fundamentals

The original plant was designed by Linde with process licensing from ICI. Linde performed the basic engineering for the loop modification and the detailed engineering for the new Isothermal reactor.  Based on the data for the new design conditions, a debottlenecking study on the distillation section was carried out in-house by our Technical Services department.  Major pre-fabrication work and in-plant erection of the loops which were to be replaced was completed before the final shutdown of the plant.  A shutdown schedule of 11 days was planned.

Outline of the Pre Commissioning Activities

The piping loops were identified and broken down into various process loops per the P & IDs. The plant was broadly classified into three independent sections: synthesis loop, makeup gas loop, and distillation loop.  This helped prioritize tasks such that the synthesis and related loops were made ready first.  The loops, which were erected before shutdown, were prepared for commissioning by flushing / blowing. Based on the service, the plans for flushing / blowing were prepared and discussed with the mechanical and instrument groups to streamline the activities. All instruments in the circuit were removed from the lines.

The following procedures were used:

For gas lines: Gasket blowing with plant air was carried out starting from 1.0 barg up to 3.5 barg repeatedly, until there was no rust / dust in the line. This was followed by nitrogen passivation / drying.

For  liquid lines: Air blowing followed by water flushing was carried out. This was followed by nitrogen passivation / drying.

For steam lines: Gradual warming of the header before insulation was applied for grease removal and rust flushing through the trap bypass.  Then steam blowing at full capacity was carried out for half an hour by diverting the open end at a safe location. The header was allowed to cool. This cycle was repeated again till clear condensate was discharged in the trap bypass.

For Running Machines: There was a pair of process pumps in each service. One pump online and one spare.  With the higher capacity, some pumps were replaced for higher capacity. The main crude feed pumps and refining column reflux pumps were replaced.  With spare pumps, the plant operation was not interrupted during the pump changes. Each replacement took 12 days and included the modification of the base, pipeline, motor, and other ancillary pieces.

Likewise, four control valves were replaced via proper coordination between the operations and project teams.  The prefabricated loops were also washed or blown and then dried with nitrogen. These were kept inert and sealed at their ends until they were to hooked up during the shutdown.  This also helped reduce the pre-commissioning time for the plant.  The start-up boiler feed water circulation pump was commissioned and stabilized prior to shutdown as soon as the errection of the reactor steam drum system was completed.

Both Methanol-I and SGGU operate independently.  It was not necessary to shutdown SGGU for the commissioning of Isothermal reactor in the Methanol-I synthesis loop.  The natural gas compressors in the SGGU plant get cooling water from the Methanol-I plant header.  Since cooling tower was to be taken offline, temporary arrangements to supply an alternate water source was planned to keep the natural gas compressors in the SGGU running.  This was implemented prior to shutdown, avoiding a stoppage of the SGGU plant.

Reactor Catalyst Charging

This was the first reactor of its kind at GNFC with spiral wound coils within the shell.  The catalyst is to be charged on the shell side, while the cooling medium (boiler feed water) flows in the tubes via thermosiphon.  During the startup, the boiler feed water circulation pump maintains the water circulation.

As shown in Figures 4 and 5, the reactor coils are supported at both the ends by six support strips moving radially outward from the central mandrill.  This divides the cross section of the reactor into six equal parts.  These were taken as the basis for charging the catalyst and numbered from 1 to 6 inside the reactor.

	Â
Figure 4: Internal View of the Isothermal Reactor	Â	Figure 5: Top Internal View of the Isothermal Reactor

{parse block="google_articles"}Two catalyst charging nozzles were used with hoppers and 2 ½” dia flexible hoses for charging the catalyst - SACK WISE under the supervision of Linde.  A table was prepared to log the number of bags charged per round and the subsequent dip achieved, which showed the packing uniformity.  This proved to be a very successful method of charging with good packing density with less than 20 mm of variation in the final height adjustment.  A flat and heavy plumb with strong cotton thread was used for taking the dip.

Approximately equal quantities of 20 kg balls/catalyst were filled in HDPE sacks before the start of loading.  About 5.2 m³ alumina balls were filled first in four rounds of sack charging.  The catalyst bed was leveled so that the balls were just inside the tube coiled bundle.  The first dip of catalyst was taken after charging almost half of the catalyst. Thereafter, while monitoring the height, charging continued to completion over approximately two (2) days.

Commissioning Activities

The synthesis loop was made available earlier than the distillation loop (6 ½ days) while the total shutdown period was compressed to 8 days by effective identification of the priority of each job.  The effectiveness of the pre-commissioning activities was evident during post-commissioning.  There were no plugged strainers, control valves by-passing, nor false signals during or after the startup of the plant. Re-commissioning of the plant was completed in less than 4 days time.  While, the distillation section modifications were being completed, the synthesis loop pre-commissioning activities were completed.  While the catalyst heat up and reduction was proceeding, crude methanol production was coming online.                       

Peak production levels for the plant were achieved while testing the plant at different feed gas mixtures. The plant has met all process guarantees. Of particular interest has been an improved yield of methanol due to a higher conversion rate and stable reaction conditions.  Less by-product formation has led to a reduction of loading in the distillation section.

Conclusions

From this experience, we see that the plant capacity can be increased by understanding the basic principles of reaction kinetics and unit operations.  Through integration of technology and the use of improved catalyst, this little plant had been transformed into a giant producer.  Proper planning of critical activities like catalyst charging, pre-commissioning of loops, commissioning and guarantee test runs can ensure success.

Styrene Monomer Production

The energy needed for the reaction is supplied by superheated steam (at about 720 °C) that is injected into a vertically mounted fixed bed catalytic reactor with vaporized ethylbenzene.  The catalyst is iron oxide based and contains Cr₂O₃ and a potassium compound (KOH or K₂CO₃) which act as reaction promoters.  {parse block="google_articles"}Typically, 2.5-3 kg steam are required for each kilogram of ethylbenzene to ensure sufficiently high temperatures throughout the reactor.   The superheated steam supplies the necessary reaction temperature of 550-620 ^°C throughout the reactor.  Ethylbenzene conversion is typically 60-65%.  Styrene selectivity is greater than 90%.  The three significant byproducts are toluene, benzene, and hydrogen.

After the reaction, the products are cooled rapidly (perhaps even quenched) to prevent polymerization.  The product stream (containing styrene, toulene, benzene, and unreacted ethylbenzene) is fractionally condensed after the hydrogen is flashed from the stream.  The hydrogen from the reaction is used as fuel to heat the steam (boiler fuel).  After adding a polymerization inhibitor (usually a phenol), the styrene is vacuum distilled in a series of four columns (often times packed columns) to reach the required 99.8% purity.  The separation is difficult due to the similar boiling points of styrene and ethylbenzene.   Typical capacity per plant ranges from 70,000 to 100,000 metric tonnes per year in each reactor and most plants contain multiple reactors or units.

Polystyrene Production

In 1996, world production capacity for styrene was near 19.2 million metric tonnes per year.  Dow Chemical is the world's largest producer with a total capacity of 1.8 million metric tonnes in the USA, Canada, and Europe (1996 figures).  The main manufacturing route to styrene is the direct catalytic dehydrogenation of ethylbenzene (above).

The reaction shown above has a heat of reaction of -121 KJ/mol (endothermic).  Nearly 65% of all styrene is used to produce polystyrene.

The overall reaction describing the styrene polymerization is:

This reaction is carried out in an inert organic solvent environment which provides the reaction medium for this cationic polymerization reaction.  The most common solvent used for this reaction is 1,2-dichloroethane (EDC).  Other suitable solvents may include carbon tetrachloride, ethyl chloride, methylene dichloride, benzene, toluene, ethylbenzene, or chlorobenzene.  The preferred initiator is a mixture of boron trifluoride and water.

The initiator solution is prepared by incorporating 1.5% by weight boron trifluoride gas into the organic solvent (EDC) containing 280 ppm water.  This solution is continuously prepared in a holding vessel and will then be injected into the reactor system.

Figure 1: Block Diagrame for Polystyrene Process

Typical feed to the first reactor would consist of 50 weight percent styrene monomer, 100 ppm water (based on styrene weight), 2000 ppm boron trifluoride (based on styrene weight), with the balance being organic solvent.   The polymerization reaction gives off heat that is carried away from the reactors by jacketing them with a heat transfer fluid.  The temperature of the reactants should not vary by more than 15 ^°C throughout the reactor series.   Temperature control is very important in this reaction because as the reaction temperature increases, the average molecular weight of the polystyrene decreases.   The reaction temperature range is 40-70 ^°C.  Temperature can also be controlled by intermediate shell and tube heat exchangers.

The reaction vessels are typically elongated vessels made of stainless steel.  The initiator is introduced as shown below:

Figure 2: Typical Reactor Overview

References

1. Weissermel, K., Industrial Organic Chemistry, 3rd Edition, VCH, New York, 19972. 
2. US Patent #4161573, assigned to Dow Chemical