Ankur's_Chemical_Engg_Blog

Deleterious Effects Of Low Velocity Flow In Piping And Pipelines

Wed, 30 Jan 2019 04:30:00 +0000

Dear All,

Most of us know that high velocities in pipelines can lead to a plethora of operational and maintenance problems such as noise, vibrations, erosion of pipe material, and a combined erosion-corrosion problem. API RP 14E (now a defunct document), defines a formula for calculating erosional velocities in 2-phase flows mentioning that actual flowing velocities should be well below the calculated erosional velocities. ISO 13703 says the same about erosional velocities since it is based on API RP 14E, with the difference that the erosional velocity formula of API RP 14E is available in SI units. My blog entry “Erosion due to Flow” wherein the subject of erosional velocity is discussed in detail, has been widely read and commented. The link for this old entry of mine is given below:

https://www.cheresources.com/invision/blog/4/entry-190-erosion-due-to-flow/

Today’s topic however deals with low velocities. The adjective I have used is deleterious for low velocities. The reason being that, while high velocities can have an immediate impact in terms of noise or vibrations in piping systems, the undesirable effect of low velocities in piping is more subtle and long-term. The adjective deleterious signifies this subtle and long-term effect of low velocities.

Let us come to those deleterious effects of low velocities in piping and pipelines. I will categorize the flow in pipes and pipelines based on the type of fluid and the fluid phase.

Flow of Slurries (liquid-solid homogeneous phase):
At very low flowing velocities, phase separation will occur of the solid particles, with the higher density solid particles tending to settle down at the pipe bottom of a horizontal pipe run. The settling will be even more at bends (direction change), or where there is a reduction in the pipe diameter. The deleterious effect is that over a period of time a layer of solid builds in the pipe and pipe fittings, reducing the pipe diameter leading to a) excessive pressure drop b ) partial or full disruption of flow c) pump operating point moving towards shut-off leading to reduced pump efficiency and accelerated mechanical wear and tear of the pump and pump sealing system.

Flow of Liquids (comprising of entrained heavier liquid with a lighter liquid):
A typical example would be a hydrocarbon liquid with entrained water, where the hydrocarbon liquid is the continuous phase while the water is the discontinuous phase. At low flowing velocities the entrained discontinuous phase water will drop out from the continuous phase hydrocarbon liquid to the bottom of the pipe. Water accumulation will occur in low points of the piping over a long term. If the hydrocarbon liquid has even trace amounts of dissolved carbon dioxide or hydrogen sulfide, pipe / pipeline corrosion can occur. The mechanism of corrosion in simplistic terms is that carbon dioxide and / or hydrogen sulfide will react with the accumulated water in the pipeline leading to acid corrosion by formation of Carbonic Acid (H₂CO₃) and / or Sulfuric Acid (H₂SO₄). Such corrosion can lead to failure of carbon steel piping / pipeline over a long term.

Single Phase Gas Flow (with entrained liquids):
A typical example would be natural gas containing entrained liquid droplets of water and heavier hydrocarbons. At low flowing velocities the flow would be stratified in the horizontal pipe, where the gas travels at the top of the pipe and the liquid travels at the bottom of the pipe and liquid accumulation occurs at low points and direction changes in pipe over a long period of time. The same problem of corrosion can occur as discussed for flow of liquids. Additionally, the gas transport could see high pressure drops and reduction in flow, putting excessive loads on gas compressors. Liquid accumulation over a long term in pipelines due to low velocities can also lead to intermittent slug flow in pipelines. The high momentum of liquid slugs can lead to structural damage of piping / pipelines and their supports.

Often natural gas pipelines have been found to have black powdery material (solids) in small amounts. The black powder could be because of corrosion products, trace amounts of solids carried over from gas treatment plants, mill scale etc. Low flowing velocities in gas transmission pipelines can lead to accumulation and deposition on pipeline walls of the black powder over a long-term and lead to excessive pressure drop and reduced flow. Flow velocities need to be kept above a threshold velocity also known as minimum entrainment velocity to prevent black powder deposits.

3-phase flow (Gas-Liquid-Liquid with Oil as continuous phase):
A typical example would be crude oil from reservoirs with associated gas (dissolved or free) and free water. Low flowing velocities will cause similar problems as discussed above related to corrosion. Additionally, if the crude oil is heavy crude oil containing asphaltenes, then reduced flow velocities (reduced flow) for a given pipeline will drastically increase the asphaltene deposition rate on the pipe walls. The long-term effect could be partial or total flow stoppage requiring costly cleaning operations for full restoration of pipeline operations.

Quantification of Minimum Velocities:
I had prepared a standard “Specification for Process Design Basis” for a middle-east oil & gas operating company wherein I had mentioned also about minimum velocities in pipelines. For unlined carbon steel pipelines transporting liquid hydrocarbons (light crude oil or condensate) containing free or entrained water even in a small quantity (e.g. 1% water cut), the velocities should not be allowed to fall below 1.5 m/s to prevent water dropout. For gas pipelines I had mentioned that the normal range of flow velocities should be 5 to 10 m/s. For bone dry gas, velocities up to 20 m/s may be allowed, subject to design considerations for noise and vibration prevention during all operational scenarios.

As mentioned earlier, black powder in natural gas pipelines will not be entrained at low gas velocities. This may lead to accumulation at some portion of the gas pipeline over a long duration. The entrainment velocity for these very fine particles is a function of their micron size. Generally, for a particle size of 1 micron to remain entrained, the gas velocity should be in the range of 2.5 to 4.5 m/s, depending on the pipe size.

Measures to prevent low velocities:
In intermittent or batch transfer operations, when demand is low at the receiving station, continue pumping at the same high rate as required for peak demand but for a shorter duration. There is no need to reduce the flow rate for low demand in batch transfer operations. Standard Operating Procedures (SOPs) should address such turndown or low demand scenario without reducing the flow and thus the velocity.

Measures to mitigate low velocities:
In continuous operations and prolonged turndown scenarios, low flow velocities are unavoidable. To prevent the aforementioned problems associated with low flow velocities, some measures that could be implemented are described below:
1. Injection of anti-corrosion and anti-scale additives in the piping / pipeline system.
2. Addition of emulsifying agents in liquid hydrocarbon-water systems to prevent phase separation occurring at low velocities.

Note: The addition of additives to prevent corrosion, deposits and phase separation should be carefully evaluated from the viewpoint of chemical compatibility with the process fluid and any adverse effects in downstream applications.

To conclude, low flowing velocities in pipes and pipelines can create problems such as corrosion, scaling and deposit formation and design and operational measures should be considered to prevent operations at low velocities.

I look forward to comments from members of “Cheresources”.

Regards,
Ankur.

Production Of Triple Superphosphate

Wed, 26 Dec 2018 13:30:00 +0000

Dear All,

Today's blog entry is in continuation of the blog entry I had made earlier for production of phosphoric acid using the wet process. The phosphoric acid produced by the wet process can be used as a raw material for the production of triple superphosphate.

The process description for Triple Superphosphate is referenced from Kirk-Othmer Encylopedia of Chemical Technology. the material balance is provided as an excel attachment to the blog entry.

Triple (Concentrated) Superphosphate. The first important use of phosphoric acid in fertilizer processing was in the production of triple superphosphate (TSP), sometimes called concentrated superphosphate. Basically, the production process for this material is the same as that for normal superphosphate, except that the reactants are phosphate rock and phosphoric acid instead of phosphate rock and sulfuric acid. The phosphoric acid, like sulfuric acid, solubilizes the rock and, in addition, contributes its own content of soluble phosphorus. The result is triple superphosphate of 45- 47% P₂O₅ content as compared to 16- 20% P₂O₅ in normal superphosphate. Although triple superphosphate has been known almost as long as normal
superphosphate, it did not reach commercial importance until. the late 1940s, when commercial supply of acid became available.

Simplicity of production, high analysis, and excellent agronomic quality are reasons for the sustained high production and consumption of TSP. A contributing factor is that manufacture of the triple superphosphate has been an outlet for so-called sludge acid, the highly impure phosphoric acid obtained as a by-product of normal acid purification.

Chemistry and Properties: TSP is essentially impure monocalcium phosphate monohydrate, Ca(H₂PO₄)₂ ·H20, made by acidulating phosphate rock with phosphoric acid according to:

The complete chemistry of TSP production has been studied and reported in great detail. As in the production of NSP also known as Single Superphosphate (SSP) there are also reactions with impurity minerals. The range of constituents in commercial TSP from wet-process acid and phosphate rocks are typically: Ca(H₂PO₄ )₂.H2O, 63-73%; CaSO₄, 3- 6%; CaHPO₄ and Fe and Al phosphates, 13- 18%; silica, fluoro silicates, unreacted rock, and organic matter, 5-10%; and free moisture, 1-4%. The average citrate solubility of the P₂O₅ in best quality TSP is 98-99%, but products with citrate solubility values a few percentage points lower are not uncommon. The P₂0₅ citrate solubility of TSP made by a quick-cure process is a ca 96%. Other common properties of TSP are bulk density, non-granular, 879 kg/m³; granular, 1040-1200 kg/m³; and critical relative humidity at 30°C, 94%.

Production Technology: Phosphate rock and wet-process phosphoric acid are the only raw materials required for manufacturing TSP. The grade of rock can be a little lower than that needed for NSP production. Over the years, a large number of process modifications, both batch and continuous, have been used. For the production of nongranular TSP, 52-54% P₂O₅ acid is used without dilution or heating. The P₂O₅:CaO mole ratio, including the P₂O₅ in the rock, is 0.92-0.95. The rock is ground to 70% <74.urn (200 mesh). Pile curing for a few weeks is typical, as for NSP.

Granular TSP (-6 + 16 mesh (1.19 to 3.35 mm dia)) is preferred for direct application and is used in bulk blend fertilizers. A widely used slurry granulation process is the Dorr-Oliver process. The ground rock is mixed with 38-49% P₂O₅ acid in a series of reaction vessels. Slurry from the reaction train then is mixed with a large proportion of recycled undersize granules and crushed oversize in either a pugmill or drum granulator. The granules are dried and screened. The product-size material is about 1 part in 13, thus the recycle ratio is 12:1. Other processes involve granulation of previously prepared nongranular TSP.

Economics: In contrast to NSP, the high nutrient content of TSP makes shipment of the finished product preferable to shipping of the raw materials. Plants, therefore, are located at or near the rock source. The phosphoric acid used, and the sulfuric acid required for its manufacture, usually are produced at the site of the TSP plant. As in the case of NSF, the cost of raw materials accounts for more than 90% of the total cost. Most of this is the cost of acid.

Since about 1968, triple superphosphate has been far outdistanced by Diammonium Phosphate as the principal phosphate fertilizer, both in the United States and worldwide. However, production of triple superphosphate is expected to persist at a moderate level for two reasons: (1) at the location of a phosphoric acid-diammonium phosphate complex, production of triple superphosphate is a convenient way of using sludge acid that is too impure for diammonium phosphate production; and (2) the absence of nitrogen in triple superphosphate makes it the preferred source of phosphorus for the no-nitrogen bulk-blend fertilizers that frequently are prescribed for leguminous crops such as soy beans, alfalfa, and clover.

The above paragraphs are edited from Kirk-Othmer Encyclopedia of Chemical Technology.

Also an excel workbook is attached providing some material balance calculations for Triple Superphosphate.

Comments and observations on this blog article are welcome.

Regards,
Ankur

Attached Thumbnails

Attached Files

TSP_Production_Flowsheet_and_Calcs.xlsx (193.14KB)
downloads: 26

Gas Turbine Power Re-Rating For Site

Wed, 12 Sep 2018 19:05:00 +0000

Dear All,

Gas turbines are important in the chemical process industry. They are often used as mechanical drivers for large centrifugal and reciprocating compressors which transport very large volumes of natural gas at high pressures over long distances.

Gas turbine vendors specify gas turbine power at standard conditions which are as per the "International Standards Organization" abbreviated as ISO. This is called the ISO rating of the gas turbine and is according to the ISO document ISO 3977-2:1997 Gas Turbines -- Procurement -- Part 2: Standard reference conditions and ratings.

Following are these conditions:

Ambient Temperature = 15 deg C
Altitude = 0 m (sea level)
Ambient Pressure = 101.325 kPa (abs)
Relative Humidity = 60%

The standard conditions also specify inlet and exhaust losses as zero.

What about actual site conditions where the gas turbine needs to be installed? Gas turbines need to be re-rated for actual site ambient conditions and also considering inlet and exhaust losses. Power and Heat Rate of the GT need to be re-rated on account of actual site conditions and inlet / exhaust losses. These require the application of correction factors to the ISO power and ISO heat rate of the GT. Today's blog entry provides the procedure for re-rating (de-rating / up-rating) the ISO power for the actual site and considering inlet / exhaust losses.

Some graphs and formulas are provided below to obtain the correction factors:

Power Correction factor for Altitude at Site (CF_alt)

Power Correction factor for Inlet Loss

where:
CF_inletloss = correction factor
PDi_nlet = Inlet Pressure drop, kPa
Note: The inlet pressure drop is in the range of 1 to 1.5 kPa (GT vendor can also provide this data)

Power Correction factor for Exhaust Loss

where:
CF_exhaustloss = correction factor
PD_exhaust = Exhaust Pressure drop, kPa
Note: The exhaust pressure drop is in the range of 0.5 to 1 kPa (GT vendor can also provide this data)

Power Correction factor for Ambient Temperature at Site

where:
CF_Tamb = correction factor
T_amb = Ambient temperature at site, ⁰C
Note: The correction factor for ambient temperature should be calculated for both minimum average ambient temperature and maximum average ambient temperature of the GT installation site and applied accordingly for re-rating the gas turbine.

After calculating the correction factors as above the site re-rated gas turbine power output can be calculated as follows:

GT Site Power = GT ISO Power*CF_alt*CF_inletloss*CF_exhaustloss*CF_Tamb
where:
GT Site Power = Re-rated GT Power at actual site, kW
GT ISO Power = Vendor Specified Power, kW

Gas Turbine Heat Rate re-rating has a similar approach and I will deal with that in my next blog entry. I would welcome comments from the members of "Cheresources" on my blog entry.

Regards,
Ankur.

Attached Thumbnails

Wet Process Phosphoric Acid Production Material Balance

Sun, 12 Aug 2018 06:18:00 +0000

Dear All,

Phosphoric Acid and the derived Phosphate fertilizers play an important role in any economy which has substantial dependence on agriculture. India happens to be one such country where agriculture contributes to approximately 17% of GDP and employs somewhat more than 50% of India's employment.

Although I have never directly worked in the fertilizer sector, I find this area of chemicals quite fascinating from a chemical engineer's perspective. Lately I had been doing some studies on production of wet phosphoric acid and the superphosphates. This led to my developing a material balance spreadsheet for wet phosphoric acid production from commercially mined phosphate rock which I am attaching with this blog entry along with the reference of the literature I have used for developing this spreadsheet. The spreadsheet also includes a process flow scheme for the manufacture of wet phosphoric acid. But before I share the spreadsheet a brief description of the process is essential. Here it goes:

Wet Process for Phosphoric Acid:

Phosphoric acid is produced from fluorapatite, known as phosphate rock, 3Ca₃(PO₄)₂.CaF₂, by the addition of concentrated (93%) sulfuric acid in a series of well-stirred reactors. This results in phosphoric acid and calcium sulfate (gypsum) plus other insoluble impurities. Water is added and the gypsum is removed by filtration along with other insoluble materials (e.g. silica). Fluoride, as H₂SiF₆, is removed at a further stage by evaporation. Although the reaction takes place in stages involving calcium dihydrogenphosphate, the overall reaction can be represented as:

However, there are side reactions; for example with calcium fluoride and calcium carbonate present in the rock:

Fluorosilicilic acid is an important by-product from this.

The crystal structure of the calcium sulfate formed depends on the conditions of the reaction. At 340-350 K, the principal product is dihydrate, CaSO₄.2H₂O. At 360-380 K, the hemihydrate is produced, CaSO₄._¹/₂H_₂O.

Calcium sulfate is filtered off and the acid is then concentrated to 56% P₂O₅ using vacuum distillation.

The product from the 'wet process' acid is impure but can be used, without further purification, for phosphatic fertilizer manufacture such as triple superphophate (TSP), Monoammonium dihydrogenphosphate (MAP) and Diammonium hydrogenphosphate (DAP). Alternatively it can be evaporated further to a concentration of 70% P_₂O_₅_, a solution called superphosphoric acid which is used directly as a liquid fertilizer.

In my subsequent blog entries I will come out with details of other phosphate fertilizers such as Single Superphosphate (SSP) and Triple Superphosphate (TSP).

The excel spreadsheet for material balance for wet process phosphoric acid (dihydrate) is attached. I would welcome comments on the spreadsheet either to point out improvements or any errors.

Regards,
Ankur.

Attached Thumbnails

Attached Files

Phosphoric_Acid_Production.xlsx (110.67KB)
downloads: 177

Spray Tower Scrubber Sizing For Flue Gas Desulfurization

Wed, 02 May 2018 07:06:00 +0000

Dear All,

Today's blog entry deals with the design and sizing of a non-packed spray tower for the purpose of flue gas desulfurization.

A spray tower sizing web resource is available on the internet which has been used as a basis for the design. However this sizing routine does not provide a clear cut method for determining the gas phase mass transfer coefficient and therein comes the research I had to do to arrive at a empirical method to determine the gas-phase mass transfer coefficient based on a reference document from the EPA (Environmental Protection Agency) of the United States..

References can be referred to from the attachments below:

The gas-phase mass transfer coefficient K_Ga has been referred from the second attachment (refer page 37 for the equation). This equation is a direct method to calculate the gas-phase mass transfer compared to the method utilized in the first attachment where this value is obtained in an indirect manner by comparing with an existing spray tower with given dimensions, inlet / outlet SO₂ concentrations and lime circulation rate.

If you compare the values of the gas-phase mass transfer coefficient K_Ga calculated from the indirect method in the first attachment and the empirical equation from the second attachment (EPA report) they are quite close based on a lime circulation rate of 60 m³/h as considered in the first attachment.

Combining the sizing methodology of the first attachment with the gas-phase mass transfer coefficient equation from the second attachment, I have developed an excel workbook which sizes a spray tower. The circulation rate for the limestone solution is considered as 3 liters per actual cubic meter of flue gas which is again referenced from another EPA report. The value of the gas-phase mass transfer coefficient is in turn dependent on the circulation rate through the spray tower. Higher the circulation rate, higher the value of the gas-phase mass transfer coefficient. The reference details are provided in the excel workbook itself.

The excel workbook can be downloaded from the file library with a link from the blog entry here:

https://www.cheresources.com/invision/files/file/330-spray-tower-for-flue-gas-scrubbing-design/

I would appreciate comments from members of "Cheresources" on my blog entry and the excel workbook.

Regards,
Ankur

Attached Files

88288739-Process-Design-of-Spray-Chamber-or-Spray-Tower-Type-Absorber.docx (476.25KB)
downloads: 439
9101DWOO.PDF_Dockey=9101DWOO.PDF (3.31MB)
downloads: 482

Product Blending - In-Line Blending From Tankage

Wed, 18 Apr 2018 05:27:00 +0000

Dear All,

It has been a good 6 months since my last blog entry. Just couldn't find the time and the right topic to post on my blog. But, I am back with a new topic. Let me have your comments and advice on this latest one.

Today's topic is about in-line product blending using sophisticated microprocessor controls and ratio algorithms. Many leading companies in industrial control and automation such as Emerson, Yokogawa, ABB can provide customized control solutions for product blending. Each of them can also provide consulting services for a optimized control solution for a given product blending application based on their prior experience.

My blog entry provides a basic introduction to in-line product blending and an attached basic flow diagram and control scheme for in-line product blending from tanks. The attachment uses the example of gasoline blending, which is one of the most widely practiced blending operations in refinery complexes. So here it goes:

In-Line Blending from Tankage - This blending method also requires the use of both component tankage and product tankage. However, the components are pumped simultaneously under flow or flow ratio control through a common blending header. The contents of the product tank are approximately on-specification at any time during the blending operation provided the flow ratios correctly reflect the current component tank properties. An in-line blending system from tankage is shown in the attachment. Each component is ratioed either to a master component on flow control or to a total blending rate signal. Analyzers provide both component and product properties input to an advanced computer blending control system which in turns control the various component flow rates. The blend quality control application calculates the component flow ratios (or component flows) required to meet blend header product quality specifications over the entire blend. Conventional on-line analyzers, such as octane engines (for RON and MON), RVP and distillation are suitable for blend quality control. However, more recently, FT-IR analyzers have been used. From strictly a blending standpoint, a single FT-IR analyzer is capable of monitoring multiple component streams as well as the finished product to allow the control system to make on-stream corrections during the blend operation to reflect minor changes in component qualities.

This type of blending system is frequently used for gasoline.

The following factors should be kept in mind when one is considering in-line blending from tankage:

1. Component tankage permits mixing and testing of components prior to blending. A minimum of two tanks per component is preferred. One tank is assigned for the blend and the other for production rundown. The tank assigned for the blend should be well mixed such that the component quality will not vary over the blend cycle. Provided component tank(s) are well mixed and isolated from production rundown, component qualities can be certified by lab analysis and periodic or continuous sampling is not required during blend operation. If blend component quality varies during blend operation (i.e. tank not well mixed and/or blend tank is not segregated from production rundown) on-line monitoring of component quality should be considered.

2. This type system is very flexible and permits minimizing product quality giveaway if components of appropriate quality are available.

3. Multiple grades per product and seasonal blend variations can often be handled sequentially through one blending header, thereby minimizing blender investment.

4. Tankage requirements to cover working and turnaround requirements may be split between blendstock
components and finished product tankage. Consistent with the concept of minimizing total tankage, component storage is preferred to finished storage since it provides better flexibility to blend multiple finished products.

5. With accurate component flow/ratio control, the blend may be considered "on-spec” as soon as all components are mixed in the blending header. Thus, while the blending is in progress, sampling of this header will permit blend evaluation and fine point blend corrections to optimize the blend. Furthermore, this technique will result in multiple sampling and testing leading to reduced standard deviation for the blend which, in turn, results in much greater test accuracy/reliability and further blend optimization.

Refer the attachment for the flow scheme and control system for the in-line blending system. Some advantages are also mentioned for an in-line blending system in the attachment.

Look forward to comments on the blog entry.

Regards,
Ankur.

Attached Files

Gasoline_Blending_PFD.xlsx (50.8KB)
downloads: 202

Steam Tracing Design

Sat, 04 Nov 2017 18:04:00 +0000

Dear All,

It is well known that chemical process plants located in areas where ambient temperatures can fall well below freezing may require heat tracing of lines and equipment containing process fluids that can freeze or congeal and stop flowing thereby causing costly plant outages and economic losses. In certain cases, freezing of line contents and flow blockage (total or partial) can result in accidents, such as rupture and release of hazardous chemicals from piping or equipment, endangering human life and the surrounding environment.

Well designed heat tracing systems will ensure that the possibility of freezing and blockage is minimized or eliminated. Universally two methods of heat tracing are practiced, electric or steam tracing. In many chemical process plants you will find a combination of both electric and steam tracing depending on the tracing application. In general, both specific application and the economics of electric or steam tracing will govern the choice of heat tracing. The design engineer will have to study both the applicability and the CAPEX / OPEX to decide the type of heat tracing suitable for the given system.

Today's blog entry specifically deals with heat tracing of using steam. An excel workbook is attached addressing the following:

1. Steam Tracer Sizing Tables: Given the process piping size, the selected tracer material (copper or CS) and the temperature to be maintained in the pipe to prevent freezing, these tables provide the number of tracers, recommended tracer tubing / piping size and the recommended steam flow rates per 100 m of tracer length.

2. Tables for Maximum Tracer length between Traps: These tables provide the maximum length of tracer, at the end of which a steam trap is needed, given the steam flow rate per 100 m of tracer length, and the steam inlet pressure to the tracer.

3. Table for maximum steam flow rate for different steam pressure levels and corresponding steam header size, based on economic pipe sizes, and for steam tracing application.

I am hoping that this workbook for steam tracer sizing proves useful to engineers who wish to design steam tracing systems. I look forward to comments from the members of "Cheresources". That is all for today.

Regards,
Ankur.

Attached Files

Steam_Tracing_Design.xlsx (42.78KB)
downloads: 628

The Perennial Debate On Piping / Pipeline Design Pressure

Tue, 17 Oct 2017 04:19:00 +0000

Dear All,

I have lost count on how many times I have debated and argued the topic of piping / pipeline design pressure. “Cheresources” itself has dozens of posts related to this topic. I am now a firm believer that most problems do not have one unique solution in context of process design. You could have multiple design approaches each with its own merit and demerit and the process engineer needs to adopt the design solution as per the demand and need of the end-user. However, whenever a process design engineer puts forward a design solution he or she should know all the pros and cons of the solution he or she is proposing, and assess its suitability for the end-user, or plant operator from all aspects such as safety, operational ease, design life and the economics of the proposed solution.

Let us move on to the topic of piping / pipeline design pressure. There are two major approaches in deciding or finalizing the design pressure.

1. Shut-off based design pressure for centrifugal machines. This is specifically true for centrifugal pumps. For centrifugal compressors, this is somewhat tricky, since you need to determine the design pressure based on surge limit curve on the overall compressor curve. A simplified approach in determining the design pressure at the preliminary design stage would be to add up the maximum suction pressure of the centrifugal pump or compressor with the normal differential pressure at the rated flow times some margin (1.25 or 1.3). This approach of determining the design pressure for piping in discharge of centrifugal machines is commonly employed in design practices of many engineering or operating companies. This normally works fine for centrifugal machines.

How about positive displacement machines? These are constant volumetric flow machines and even if you block the discharge the machine will keep on pushing a constant volume in the connected discharge pipe. The consequence would be a rise in pressure with blocked outlet eventually leading to the motor driver of the PD machine stalling and / or discharge piping damaged (ruptured / distorted). As such, there is no real concept of shut-off head or pressure in PD machines since the pressure will keep on building up till either the motor stalls and / or the pipe or flange gasket ruptures. Thus, the approach adopted in the aforementioned paragraph for design pressure of centrifugal machines is not likely to work in case of PD machines.

PD machines discharge systems are protected by pressure safety valves (PSVs) set at such a pressure that it will protect the entire discharge piping system including pipe and other components in the discharge system from damaging overpressure. Providing the set-pressure for the PSV requires a different approach. There could be multiple PSVs protecting various sections of the discharge system with a common design pressure and set at that pressure. If equipment in the discharge system has a lower design pressure rating, PSV could be provided for that individual equipment with its set pressure equal to the equipment design pressure. Another approach could be to fully rate the discharge system, which means that the entire discharge system including piping and connected equipment would have the same design pressure throughout the system. Fully rating the system and minimizing block or isolation valves in the discharge system would allow to provide minimum number of pressure relief devices.

It is important to note that, if a blocked-in scenario between two isolation or block valves can be justified for a section of the liquid full discharge system, it would be essential to provide a pressure relief device for thermal expansion for that section. The key to this statement is liquid full. It does not apply for gas filled pipe sections that can be isolated.

2. Design pressure based on pipe pound rating class and using the pressure-temperature rating as provided in ASME B16.5 (formerly ANSI B16.5). For new engineers ASME B16.5 is the American standard for “Pipe Flanges and Flanged Fittings”. Although the standard is an American standard it has found universal acceptance in piping design and engineering. The basic premise of deciding the design pressure as per ASME B16.5 is that a flange is the weakest joint in any piping system and this is universally true, that the strength of any system depends on the weakest component in the system. I personally find this a more appropriate approach for assigning a design pressure value of any piping system based on a given design temperature.

ASME B16.5 has extensive tables for pressure rating against a given temperature for various pipe pound ratings and various pipe materials of construction. Pound ratings start from 150# and go up to 2500#. Material classes are designated with numbers starting from Material Class 1.10 to 1.18, Material Class 2.1 to 2.12, and Material Class 3.1 to 3.19. Interpolation can be done for temperature values between two temperature values given in the tables.

Let us take an example:

Pipe pound rating: 300#
Material Class: 1.1 (This material class in general defines carbon steel material)
Design temperature: 80°C
Design Pressure: 47.6 barg (calculated from ASME B16.5)

The Pressure-temperature rating for Material Class 1.1 in B16.5 is the first table. It gives a pressure rating of 50.1 barg corresponding to a temperature rating of 50°C and a pressure rating of 46.6 barg corresponding to a temperature of 100°C. If we interpolate for 80°C, the corresponding pressure would be 47.6 barg.

If I must provide a PSV on the 300# piping with material class 1.1 for blocked-in conditions and hydraulic expansion of trapped liquid in this pipe or pipeline, commonly referred to as a “Thermal Relief Valve” or “Thermal Expansion Relief Valve”, I would give it a set-point of 47.6 barg.

This approach for assigning a design pressure would be good even for a centrifugal pump because it provides you flexibility in operations. For example, most centrifugal pump manufacturers would provide an option of 3 different impellers with different flow versus head curves for the 3 impellers. If for operational flexibility reasons, an impeller is changed to have a higher differential head, then any pressure relief devices which were originally provided with a set pressure based on design pressure calculated based on shut-off pressure for the rated impeller, would be of no use and would require the PSV to be reset for a higher set pressure for the changed (higher differential head) conditions.

I consider it to be a good engineering practice to fully rate the pump discharge piping system (including pipe, fittings and equipment) which means that the design pressure is the same throughout the piping / pipeline from the pump discharge. There may be arguments about the high cost of having a fully rated system, but it is important to remember that any failure has its own cost in terms of personnel safety, environmental hazards and lost production which in most cases would be higher than having a fully rated system. Not having a fully rated system would also require detailed failure mode and effects analysis study to determine what additional safeguards are required to protect the integrity of the system.

Although I have categorically mentioned my preference for fully rated piping / pipeline systems, the industry practice for oil / gas flow lines from well heads is not to fully rate them downstream of the choke valve. The cost of fully rating the long distance flow line based on choke valve upstream would be astronomical. SIL rated “High Integrity Pressure Protection Systems” (HIPPS) are employed to protect the flow lines downstream of the well head choke valve in case high pressures are encountered due to choke valve failure. API 14C provides recommendations for providing safety systems in upstream oil / gas installations including well head flow lines.

I have discussed enough on this topic for today and the it is open for discussions and comments from members of “Cheresources”.

Regards,
Ankur.

Thermal Profiling Of Long Distance Pipelines Carrying Chemicals (Gas And Liquid)

Sun, 10 Sep 2017 09:27:00 +0000

Dear All,

One of the important tasks while doing a pipeline process study is to do a thermal profile study of a pipeline for various operating scenarios and ambient conditions. However, what are the practical aspects of doing a thermal profile are mostly unknown to majority of new engineers who undertake this exercise. In fact many young engineers are not even aware that a thermal profile for the pipeline needs to be developed to understand the problems that can occur in pipeline transport.

Let me first list down some problems related to thermal effects in a long distance pipeline:

1. For a gas pipeline frictional pressure drop in the pipeline can lead to Joule-thomson cooling in the pipeline which is an isenthalpic process. It is possible that flowing gas temperatures over a certain pipeline length may fall below the hydrate formation temperature of the pipeline gas, which can lead to ice crystal formation in the pipeline, which in turn can lead to pipeline flow disruption by partial or full pipeline blockage. In a worst case scenario, this can lead to a pipeline rupture with loss of property and human lives. Gas hydrate formation is favored by high pipeline pressures, and already low transportation temperatures such as in deep sub-sea pipelines. Thus a thermal study becomes absolutely essential to determine whether such an event can occur and what mitigative measures need to be taken.

2. Depressurization of gas pipelines during emergency or routine maintenance operations is another issue which needs attention in terms of thermal study of the depressurization operation. Adiabatic depressurization can lead to extremely low temperatures at the depressurization source, and immediately downstream of the depressurizing device (valve or orifice). Ordinary carbon steels can suffer catastrophic brittle fracture at very low temperatures compromising pipeline safety. Normal commercial carbon steel grades often used for fabricating pipe such as ASTM A 106 and ASTM A 53 are not designed for temperatures below -29 deg C. In process engineering terms the pipe needs to be categorized for its "Minimum Design Metal Temperature" abbreviated as MDMT. Among carbon steels ASTM A333 Gr. 1 and Gr.6 are suitable for low temperature service up to -45 deg C and are impact tested for that temperature.

A thermal study of the depressurization process can provide the lowest temperatures that a pipeline can see, and thus define the MDMT of the pipe material. As an example, if a depressurization thermal study indicates that the temperature of pipeline can fall below -29 deg C, the material specialist is most likely to suggest to go for "Low Temperature Carbon Steel" abbreviated as LTCS along with a brittle test for the pipe material known as "Charpy V-Notch Impact Test" to determine the brittleness of the material at low temperatures. If the depressurization thermal study indicates temperatures falling below -46 deg C, and up to -80 deg C, thus further lowering the MDMT of the pipe material, then the material specialist may further require to upgrade the metallurgy to Duplex Stainless Steel (DSS). Cryogenic temperatures either during normal operation or maintenance tasks, may require selection of austenitic stainless steels. Again this would be determined by thermal studies of the process system.

3. Pipelines transporting crude oils, specifically highly viscous and with high asphaltene content, require a thermal profiling to determine whether the pipeline temperatures fall below the pour point or not along the pipeline length. Possible pipeline blockages due to temperatures falling below pour point can lead to production disruption, expensive cleaning, dechoking operations, and heavy economic losses. Based on the thermal profiling and determination of lowest temperatures in the pipeline, remedial measures can be taken to prevent such an occurrence.

4. Transportation of LPG (propane and propane:butane mixtures) in pipelines is another area which requires particular attentions in terms of thermal studies of the pipeline. Exposure to high ambient temperatures for above ground pipelines or sections of above ground pipelines of LPG require a careful thermal study, to ensure that the LPG does not vaporize in the pipeline at the operating pressure,which can lead to vapor locking and 2-phase flow at the determined temperature. At the destination end of the pipeline, if there is a flow or pressure control valve, and due to temperature rise of the LPG along the pipeline length there is partial vaporization at the inlet of the control valve leading to inlet 2-phase flow, it is very detrimental for the operation of the control valve in terms of valve integrity. In a nutshell, it is absolutely undesirable to have 2-phase flow in a LPG pipeline, and a thermal study would indicate, whether such a phenomena occurs across the length of the pipeline from the source to the destination.

However, when doing such thermal studies for above ground pipelines exposed to the ambient atmosphere it is absolutely essential that realistic ambient conditions are used as inputs. To put this in perspective, a good designer would consider average high or low ambient temperature conditions and average high or low wind speeds over a time duration (say summer / winter or 1-year average conditions) rather than one-off high or low temperature and wind condition values which are highly unlikely to re-occur. To put it somewhat crudely, freak weather values of temperature and wind velocities should not be used for thermal profiling of pipelines exposed to the ambient. Using such freak ambient conditions may have a huge economic cost, which may not be justified for normal pipeline operating conditions. Abnormal ambient conditions would actually dictate that pipeline operation be stopped or curtailed.

5. There are many pipeline and process simulation software that can perform thermal studies / profiling of pipelines. Obviously, software dedicated for pipeline simulation such as PIPESIM, PIPEPHASE, PIPELINE STUDIO, AFT, OLGA would be the best bet for performing such thermal studies. But general process simulation software such as Aspen HYSYS, UniSim, Aspen Plus are also capable of performing thermal profiling of pipelines with certain limitations. Aspen HYSYS has a very nice "Depressurization" utility which can provide thermal profiles of any equipment or pipeline being depressurized. The key to using any such simulation software is that you exactly know what you need from the thermal study and provide the right inputs including checking the veracity of the default values the software uses for your particular study, and if required, to change those default values to suit your particular case study. In a nutshell, use the software as a process or chemical engineer and not as a data entry operator.

This is all for today folks. This has been a long blog entry but I hope you enjoy what I share with you and I look forward to your comments and observations.

Regards,
Ankur.

Actual Orifice Areas For Different Relief Valve Manufacturers Based On Api 526 Orifice Designations

Tue, 29 Aug 2017 17:11:00 +0000

Dear All,

API 526 provides provides orifice designations such as D, E, F..... etc. for relief valves with corresponding orifice cross-sectional areas. However, most manufacturers of relief valves provide the same API 526 designations with different areas (higher than as given in API 526) using a de-rated coefficient of discharge for the orifice. They call it the "actual orifice area" or "actual discharge area" which when de-rated with their own coefficient of discharge becomes equal to the API 526 orifice area.

In practice, if you are calculating the flow rate for a relief valve and specifically for the purpose of inlet / outlet line sizing and for fulfilling the inlet pressure drop criteria of 3% of set pressure and outlet line pressure drop of 10% of set pressure (conventional relief valve) then you must use the "actual orifice area" in the API 520 Part 1 formulas for flow rate for a given orifice as per the specific manufacturer.

Today's blog entry provides an excel table for the various reputed relief valve manufacturers across the globe whose "actual orifice area" is higher than the area as given in API 526 corresponding to the particular alphabetical orifice designation. This table should be useful if somebody selects a specific model of a given manufacturer such as "Farris" or "Leser". Not all manufacturers or models are covered but some more famous manufacturers and their common models are covered for the benefit of the members and readers of "Cheresources" and they can use it to determine the actual flow rates using formulas of API 520 Part 1.

It is important to note that this excel workbook does not include orifice areas as given in ASME Section VIII and covers only API 526.

Members and readers are welcome to provide their comments on the blog entry and the attached excel workbook.

Regards,
Ankur.

Attached Files

Actual_Orifice_Areas_API526_Type_PSVs_Vendor_Data.xlsx (12.36KB)
downloads: 494

Centrifugal Compressor Surge Control Schemes And Control Elements

Sun, 09 Jul 2017 06:26:00 +0000

Dear All,

I have written before on centrifugal compressor surge and anti-surge control. Refer my blog entry at the following link:

https://www.cheresources.com/invision/blog/4/entry-211-compressor-surge-and-anti-surge-control/

The above mentioned blog entry also provides an excel workbook for basic anti-surge control

In continuation of my further studies on surge control I have now prepared another excel workbook which provides variations in surge control schemes and also provides requirements and recommendations for the control elements used in centrifugal compressor surge control.

The basic source for these schemes and recommendations related to surge control is:
APPLICATION GUIDELINE FOR CENTRIFUGAL COMPRESSOR SURGE CONTROL SYSTEMS - Release Version 4.3 prepared as a guidance practice by Gas Machinery Research Council (Southwest Research Institute)

Some other open articles on the internet and company standards have also been referred to make the excel workbook more comprehensive.

The effort has been to provide information through bullet points against each item for ease of understanding.

I would appreciate comments from the esteemed members of Cheresources to enhance the workbook in terms of additional information and quality of information, specifically with respect to modern methods and tools for centrifugal compressor surge control.

The workbook is attached.

Regards,
Ankur

Attached Files

Centrifugal_Compressor_Surge_Control_Systems.xlsx (130.62KB)
downloads: 365

Descriptive Comparison Between A Centrifugal And A Reciprocating Compressor

Sat, 03 Jun 2017 03:29:00 +0000

Dear All,

It has been really a long time since I have posted a new entry. I was preoccupied with work and also did not find the right topic to post.

My fascination with compressors continues. I was reading some articles on compressors and this one particular article struck me to be quite interesting. The article gives a descriptive comparison between a centrifugal and reciprocating compressor. The article is titled as: "What's Correct for My Application - A Centrifugal or Reciprocating Compressor" authored by very distinguished compression equipment specialists: Paul Gallick - Senior Applications Engineer, Elliott Company; Greg Phillipi and Benjamin F. Williams of Ariel Corporation. The link for the article is given below:

http://turbolab.tamu.edu/proc/turboproc/T35/15-GALLICK.pdf

The article provides descriptive comparison of a centrifugal compressor and reciprocating compressor on various parameters such as flow, pressure, efficiency, discharge temperature, materials of construction, cost etc.

I have presented the contents of the aforementioned article in an excel workbook template with minor changes here and there, based on various other sources and my own understanding of compressors. The idea was to make this comparison more compact and easier on the eye for engineers who would be interested to enhance their understanding of compressors.

The excel workbook is attached with this blog entry. I look forward to comments on the compilation I have done.

Regards,
Ankur

Attached Files

Comparison_Centrifugal_vs_Reciprocating_Compressor.xlsx (42.8KB)
downloads: 772

Pipeline Pig Launchers And Pig Receivers - Design Codes

Sat, 19 Nov 2016 12:05:00 +0000

Dear All,

After a long hiatus I am back on my blog. That the blog has crossed a milestone of 1.5 million views is a huge encouragement and I wish to thank the readers of my blog for such an overwhelming response.

Pig Launchers and Pig Receivers have been discussed many times on "Cheresources". However, I understood that there exists a confusion regarding the design codes to be applied for pig launchers / receivers. Today's blog entry tries to explain the design codes related to them and some rationale for them.

Older designs of pig launchers / receivers were based on the ASME Section VIII Div. 1 pressure vessel code. While these designs still exist and work, the modern design approach is to design them according to the connected pipeline or pipe code.

For fresh engineers new to piping / pipeline standards and codes, the following piping / pipeline codes and standards are followed almost universally. Some countries have developed their own standards which are also mentioned below:

1. ASME B31.3: which governs the design of process piping (code) (USA)
2. ASME B31.4 which governs the design of liquid pipelines(code) (USA)
3. ASME B31.8 which governs the design of gas pipelines (code) (USA)
4. ISO 13623:2009: Petroleum and natural gas industries - Pipeline transportation systems (standard)
5. BS EN 14161:2011: Petroleum and natural gas industries - Pipeline transportation systems (MODIFIED version of ISO 13623:2009) (standard) (European)
6. ISO 13703:2000: Petroleum and natural gas industries - Design and installation of piping systems on offshore production platforms (standard)
7. PD 8010-1:2015: Steel Pipeline on Land (code) (European)
8. PD 8010-2:2015: Subsea Pipelines (code) (European)
9. CSA Z662:2015: Oil and gas pipeline systems (standard) (Canadian)

API RP 14E: Recommended Practice for Design and Installation of Offshore Production Platform Piping Systems is a 1991 recommended practice and has not undergone a revision since 1991.

Among the aforementioned codes / standards the ASME codes are quite popular for pipeline / piping design.

Hence a pipeline designed as per ASME B31.4 should have a launcher / receiver designed as per the same code and a pipeline designed as per ASME B31.8 would have launcher / receiver as per the same code. In a rare case when a pipeline is designed as per ASME B31.3, this would apply for the launcher / receiver as well.

Designing a pig launcher / receiver as per PV code ASME Section VIII Div.1 does not make economic sense. Generally, given the same pipeline design parameters, a pig trap designed from ASME B31.3 will tend to be costlier than for one designed from ASME B31.4 or B31.8 because of the difference in material grades and thicknesses. In addition, a pig trap designed from ASME Section VIII Div.1 will tend to be more expensive than one designed from ASME B31.3 for the same reasons.

Normally a pig launcher / receiver for "liquid pipelines" which is designed as per pipeline code is provided a thermal expansion relief valve (TRV). This is to ensure its protection from overpressure due to thermal expansion of trapped liquid. For gas or two-phase fluid pipelines TRVs are not required.

Designing the launcher / receiver relief valve for external fire case is impractical although theoretically possible. In practice pigging is an intermittent operation and well planned and monitored. Before and after pigging operations it is ensured that the launcher and receiver are drained / vented to prevent any residual fluid inside. If a fire effects an empty launcher / receiver, the launcher / receiver will rupture even before the PSV designed for fire case pops at the set pressure. So practically it does not make a sense to have a PSV for fire case. Refer the link below for a very enlightening discussion on PSV on launcher / receiver. Specifically the post by don1980

http://www.eng-tips.com/viewthread.cfm?qid=360185

One of the reasons that engineers consider a PSV for fire case is because the pig launcher / receiver is designed as per pressure vessel code. If it were to be designed as per pipeline code then at the most for a liquid system launcher / receiver a TRV (for thermal expansion) may be provided which does not require any sizing calculations and providing a conventional 3/4 x 1" or 1"x 1-1/2" PSV would suffice. For single-phase gas or 2-phase fluid no PSV would be required.

Quick note from the admin: You can download the MS Excel workbook that accompanies this blog entry in the File Library.

That is all for today's blog entry. Look forward to comments and observations from the readers of my blog.

Regards,
Ankur.

Side Stream Filtration Rate For Cooling Towers

Mon, 18 Jul 2016 19:48:00 +0000

Dear All,

Suspended solids (particulates or turbidity) are frequently found in cooling towers. At best these suspended solids will tend to settle out in low velocity areas of the water system (such as the cooling water basin) where they can become a breeding ground for bacteria, requiring frequent cleaning and flushing of the basin. At worst, they can degrade system heat transfer capacity and lead to steep increase in dosing rate and consequently the cost of water treatment chemicals (e.g bactericides).

Modern day cooling tower systems have thus adopted side stream filtration of the water from the cooling tower. Various schemes can be implemented. The basin water inventory may only be filtered and / or side stream filters may be employed at the discharge of the CW recirculation pumps. Refer the simplified flow scheme of side stream filtration for a cooling tower:

When the side stream filter is employed in the discharge of the cooling water recirculation pump discharge, the flow rate required for filtration through side stream filter can be calculated by the following formula:

Side Stream Flow Rate, L/s = ((ppm suspended solids in recirculating cooling water / 200) - 1)*Blowdown Rate, L/s - Metric

Side Stream Flow Rate, gpm = ((ppm suspended solids in recirculating cooling water / 200) - 1)*Blowdown Rate, gpm - USC

From the above formula, it becomes obvious that if the ppm level of suspended solids is ≤200 ppm in the recirculation cooling water, side stream filtration may not be required. However, installation of side stream filter is recommended since suspended solid concentration builds up over time and periodic high build-up would require side stream filtration for the aforementioned reasons.

Pressure sand filters with backwash facilities are commonly employed as side stream filters for cooling towers.
Backwash water for the filters can either be from another source, or if cooling water is used, this water should be considered as part of the cooling tower blowdown. In either case, backwash effluent water must be sent to the waste water treatment plant.

That is all for today's blog entry. Comments are most welcome.

Regards,
Ankur.

Attached Thumbnails

Power Recovery Turboexpanders - Shaft Power Available

Thu, 23 Jun 2016 09:39:00 +0000

Dear All,

Commercial applications of high temperature power recovery turboexpanders use the expander to drive an air compressor which charges the process unit that produces the high temperature gas that expands to atmosphere across the turboexpander. The compressor/expander unit, together with its process unit, is therefore very similar to a special purpose combustion gas turbine engine cycle, with the process unit as the "combustion chamber."

Power recovery from Catalytic Cracker flue gases is one of the most important applications using turboexpanders in the petroleum refining field. FCCU Hot Gas Expanders are commercially available in speeds from 3600 to 7500 rpm, gas inlet conditions to 4.8 bara (70 psia) and 760°C (1400 deg F). Exhaust pressure is typically 1.24 bara (18 psia), mass flows range from 200,000-640,000 kg/h (440,000-1,410,000 Ib/hr). Power recovered ranges from 2600-37,300 kW (3500-50,000 hp). Units are typically 1 or 2 stages.

Today's blog entry provides the calculation for shaft power available from a power recovery turboexpander:

where:
η_i = Isentropic Efficiency. Use 85% for "as new" efficiency until specific vendor estimates are obtained. Assume an average isentropic efficiency 5% lower than the "as new" condition for economic studies that allow for erosion during a three year run.
η_m = Use 99%, allowing for bearing losses, and assuming no gear unit.
M = molecular weight of inlet gas, kg / kg-mol (lb / lb-mol). Assume value as 30.2 till specific flue gas analysis is available.
R = gas constant 8.31447 (1545) kJ / kg-mol.K (ft-lbf / lb-mole.⁰R)
W = Mass flow rate, kg/s, (lb/min)
T₁ = Inlet temperature, K (⁰R)
T₂ = Exhaust temperature, K (⁰R)
k = specific heat ratio; use 1.3 unless the flue gas analysis gives a better value
P₁ = gas pressure at expander inlet flange, kPaa (psia)
P₂ = gas pressure at expander exhaust flange, kPaa (psia)
Z = avg compressibility factor ((Z_in + Z_out) / 2)

Expander Exhaust Temperature:

The expander exhaust temperature is used to calculate the remaining flue gas heat which is still available for recovery by other means.

That is all for today. Hope to have comments and observations from members.

Regards,
Ankur.

Attached Thumbnails

Determining Potential Natural Gas Liquids (Ngl) In A Natural Gas Stream

Sat, 04 Jun 2016 10:52:00 +0000

Dear All,

The amount of potentially recoverable NGL based from a natural gas stream can be estimated from the following formula:

Potential NGL component (tons per annum) = V*y*CF
where:
V = Volume flow rate of natural gas, MMSm³/day (Std conditions are P= 1.01325 bara, T = 288.15K)
Note: MM= Million
y = mole percent of the component in natural gas
CF = Component factor (see table below)

Example: 3 MMSCMD of NG is available for processing. This gas contains 5.4% mol propane. What is the potential propane recovery from the NG?

Calculations:
Potential Propane Recovery (w/o Recovery factor and plant on-stream factor) = 3*5.4*6800 = 110160 t/a

Considering a plant recovery factor of 70% and plant on-stream factor of 335 days, the average propane recovery per annum would be:

110160*0.7*(335/365) = 70,770 tons per annum (t/a)

That is all for today. Any comments are welcome.

Regards,
Ankur.

Attached Thumbnails

First Approximation Of Fuel Gas Consumption For Gas Turbine Driven Compressor Station

Sat, 14 May 2016 16:56:00 +0000

Dear All,

A lot of us are required to do preliminary utility estimation calculations during the Feasibility / Concept stage of a project. Another terminology often used is the Pre-FEED stage of a project. In fact this is often where non-availability of data or insufficient data becomes a bottle-neck. In such cases some minimum data with certain empirical calculation methods help to arrive at figures which may be used to for high end (+/- 30%) cost estimation of a project and help establish its feasibility.

Today's blog entry as the subject line says, is with regards to compressor stations, for transport of natural gas through long-distance pipelines. An approximation for fuel consumption can be done using certain empirical formulas. The formula is presented herewith:

F = (14 / NHV)*(P_turbine + P_auxiliary)

where:
F = Fuel Gas consumption of the compressor and the auxiliary systems associated with the compressor, Sm³/h
NHV = "Net heating Value" of the Fuel Gas to be used, MJ/Sm³ (eg: Natural Gas ~ 41 MJ/Sm³ with a certain composition)
P_turbine = Turbine power requirement, kW
P_auxiliary = Power consumption in kW of auxiliaries associated with the compressor package, which includes electricity, lubricating oil system, cooling system etc.. This has a strong dependency on local climatic conditions and on the extent of gas sourced power used for such auxiliaries.

Note: Sm³ refers to standard cubic meter at standard conditions of temperature (15 deg C ) and pressure (101.325 kPa (abs))

That is all for today folks. Look forward to comments from your side.

Regards,
Ankur

Solid-Liquid Mixing In Agitated Vessels (Just Suspended Speed)

Fri, 08 Apr 2016 11:39:00 +0000

Dear All,

Solid suspension and dispersion in a liquid is an important operation carried out in mixing systems. The primary objective of solid-liquid mixing is to create and maintain a solid-liquid slurry, and to promote and enhance the rate of mass transfer between the solid and liquid phases. Such processes are typically carried out in mechanically agitated vessels and reactors. In agitated vessels, the degree of solids suspension is generally classified into three levels: on-bottom motion, complete off-bottom suspension, and uniform suspension. Refer the sketch below. For many applications, it is often important just to provide enough agitation to completely suspend the solids off the tank bottom. Below this off-bottom particle suspension state, the total solid-liquid interfacial surface area is not completely or efficiently utilized. Therefore, it is important to be able to determine the impeller agitation speed, N_js, at which the just suspended state is achieved by the particles.

Correlation for Just Suspended Speed:

where:
N_js = minimum impeller speed to just suspended solid particles in vessel, rps
S = Zwietering constant
ν = Kinematic viscosity of the liquid, m²/s
g_c = gravitational constant = 9.81 m/s²
ρ_s = denisty of solid particle, kg/m³
ρ_l = denisty of liquid, kg/m³
X = mass ratio of suspended solids to liquid or solid loading'=kg of solids / 100 kg of liquid
D = Impeller diameter, m

Table for Zwietering Constant (S):

Snapshot of an Example Problem in Excel:

This concludes today's blog entry. Would to receive comments from the learned members of "Chereosurces"

Regards,
Ankur.

Attached Thumbnails

In-House Software Validation - An Iso 9001 Perspective

Sat, 12 Mar 2016 17:26:00 +0000

Dear All,

Many of the engineers working in engineering consulting organizations know that a lot of in-house engineering calculation templates are utilized for equipment sizing, rating and other engineering applications.Many commercial software are also utilized in such organizations involving lease-and-license agreements with suppliers and developers of such software.

But what many engineers are not aware that if your organization is ISO 9001:2008 (Quality Management Systems-Requirements) compliant or is desirous to acquire an ISO 9001 certification then part of the compliance to ISO 9001 requires that the in-house developed software is validated. What does the term "validated" mean? Today's blog entry talks about validation of in-house software.

Please note that commercial software, for example simulation software such as ASPEN PLUS, ASPEN HYSYS, HTRI, PIPESIM do not require validation by the end-user. The validation for such software is the responsibility of the developers and licensors of such software and not the end-user. Many a times the technical term "Firmware" is used for such software which means it is a permanent software programmed into a read-only memory and cannot be altered or modified by anybody except the developers of the software.

The best definition I could find for validation is as follows:

"To establish the soundness, accuracy, or legitimacy of"

In context to this blog entry it means to establish the soundness, accuracy or legitimacy of the in-house engineering calculation templates. The question then arises that how to validate these calculation templates and let me assure you that it is a very simple task. Having said that, a task is a task and needs effort and time to accomplish.

What needs to be done? Since I personally have been involved in such validation procedures, I can provide a very simple explanation on how to proceed for validating in-house software.

MS-Excel is one of the simplest and most effective calculation software tool for in-house engineering calculations and most organizations subscribe to using Excel as a tool of choice for developing engineering calculations. Many of the engineers who subscribe to "Cheresources" for knowledge sharing and guidance are well versed with calculations done in excel spreadsheets.

A good calculation template will have a logical progression in steps for calculations, wherein the input data will be defined sequentially and the outputs will be generated in terms of numerical values for the desired parameter based on certain equations and formulas. However, the standard norm for excel calculation templates is to protect the output cells which contains formulas and only allow the input cells to remain unprotected. Essentially, the excel worksheet or workbook needs to be password protected to prevent inadvertent changes in the equations and formulas leading to erroneous results. When an excel worksheet is protected the formulas do not remain visible and the user does not know the formula used unless it is separately defined in the worksheet using for example Microsoft Equation 3.0. I personally follow the practice of defining the equations to ensure complete unambiguity. Even if the formula is visible, the source or reference of all the formulas is rarely mentioned in order to make the spreadsheet less complicated and cumbersome.

The spreadsheet or workbook calculation template as prepared above by any engineer in his organization and used to generate information for clients and 3rd parties is not validated. This has to be validated if ISO 9001 compliance has to be met. Let us proceed on how to validate this calculation template in a manner where it not only is ISO 9001 compliant but makes any new engineer very comfortable to use it without having doubts about its correctness and legitimacy.

In earlier days all engineering calculations were done on paper (engineering record sheets) using a blue or black pen. These engineering record sheets had a "Prepared", "Checked" and "Approved" block on every sheet where the initials and / or signature of the preparer, checker and approver were to be printed by pen. To validate your in-house excel worksheet you need to do the same.

Use a live calculation from an existing or archived project. Write down all the inputs required for the calculations, the formulas used and their source and /or reference in the engineering record sheets. For complex arithmetic or algebraic calculations break it down into two or more steps (which is generally not required in excel spreadsheets). Attach hard paper copies of the source and / or references to the calculation record sheets as annexures. Highlight the formulas, equations or sentences / paragraphs in the annexures for sake of clarity and ease.

Once the person who has prepared the paper document is through with his task, this paper document will pass to the "Checker" for checking. After due diligent checking, corrections, if any, will be done by the engineer who has prepared the document. If corrections have been done, it will undergo a 2nd check by the checker. Once the checker clears the document, it will move to the approving authority. A diligent approving authority will do a careful scrutiny of the paper document, and on finding any errors or misses, will discuss with the preparer and the checker. If required, other specialist engineer(s) may be involved in the discussion to resolve any differences of opinion. Once a consensus is reached on the correctness of the document including an exact match of the calculated outputs with those in the excel workbook, the validation paper document becomes final, and the preparer, checker and approving authority will print their initials / signatures on the paper document. This paper document will have to be filed diligently by the department quality representative with complete traceability. Traceability means that it is a numbered (alpha-numeric or numeric) document as per the numbering system adopted in the department quality work procedure or work instructions belonging to the organization ISO 9001 QMS.

This completes the validation procedure.

In addition, a supplementary MS-Word document with may be prepared which provides step-wise English language instructions on data entry in the excel calculation template. This would be specifically useful when calculations are long and complex and require elucidation in simple English. This word document could also be attached as an annexure to the main calculation hand printed paper sheets.

What happens if I don't validate? Very simply, any ISO 9001 auditor who audits the department quality management systems, has it in his rights, to give a "Non-Conformance Report" (NCR) for a not validated in-house software used for business purposes (projects), and demand a "Corrective Action Request" (CAR) report from the Department or Company Management Representative (MR). It means, that the CAR gets closed only when the validation document is prepared, is traceable, and easily retrievable.

Too many NCRs and CARs put a question mark on the effectiveness of the quality management systems of the organization and ultimately reflect badly on the organization.

In today's blog entry I have tried to explain briefly how you can validate your in-house engineering calculation templates. I hope many of you will ask questions and initiate a debate on such an important activity.

Regards,
Ankur.

Heat Transfer In Buried Liquid Lines (Stagnant Liquid)

Sat, 13 Feb 2016 18:52:00 +0000

Dear All,

My earlier blog entry on heat transfer in buried liquid pipelines related to temperature reduction of the flowing liquid at a certain distance from the pipeline inlet.

Let us think of a buried pipeline or pipeline section where the liquid flow has been stopped and the liquid is now in a stagnant condition in the pipeline or pipeline section. Let us consider that the time at which the pipeline or pipeline section contents are completely at rest (stagnant) is zero i.e, t = 0 seconds. Now let us say that the average temperature of the stagnant liquid at time t = 0 seconds is T₀, degrees C. If T₀ is higher than the surrounding (soil at pipeline burial depth) temperature than it will start decreasing with time due to heat transfer from the stagnant liquid in the pipeline to the surrounding. We want to know the drop in temperature of the stagnant liquid after a given time period of the stagnant liquid. In other words, we are required to calculate the temperature T_t at a time interval of t = t seconds. The unit seconds are used because the equations require the unit of time as seconds. From a practical point of view a realistic temperature drop will occur in hours or days and time in hours or days needs to be converted into seconds.

Today's blog entry provides the relevant equations for a fair approximation of the temperature drop of a stagnant liquid (no flow) in a buried pipeline. These equations can be used with a fair degree of accuracy for heavy crudes being transported in buried pipelines within the constraints of the value of the empirical constant 'β' selected in the range provided. Let us go to the equations:

where:
T_S = soil temperature at pipeline depth, K
T_t, T₀ = oil temperature at time t and 0 respectively, K
F = dimensionless 'Fourier' number

where:
λ = thermal conductivity of liquid, W/(m-K) (for crude oil, λ = 0.13 W/m-K)
t = time, s
ρ = oil or liquid density, kg/m³
C_p = specific heat capacity of liquid, J/(kg-K)
d = pipeline internal diameter, m

β = empirical coefficient (ranges from 6 to 10), depending on heat transfer in the soil around the pipe, the pipe coating, pour point of the liquid and wax layer on pipe wall. A higher value of β means higher thermal conductivity of these layers.

Specific Heat Capacity of Crude Oil

where:
C_p = Specific heat capacity at temperature T₀, J/kg-K
T₀ = Temperature, ⁰C
ρ = crude oil density at temperature T₀, kg/m³

A snapshot of an example calculation in an excel workbook is shown below to give an idea of the temperature drop at time t or in other words the numerical value of T_t.

That is all for today. Would be happy to have comments from members of "Cheeresources".

Regards,
Ankur.

Attached Thumbnails

Natural Gas Teg Dehydration Unit Troubleshooting

Fri, 22 Jan 2016 07:11:00 +0000

Dear All,

Today's blog entry is dedicated to my dear departed father who passed away to his heavenly abode on January 05, 2016. He was a man of literature and whatever little writing skills I possess, it was he who inspired and encouraged me. I guess, I inherited a bit of writing skills from both of my parents, my father being a poet and short-story writer and my my mother (still an active writer at 77 years of age) a novelist and a story writer.

Any natural gas processing installation is normally bound to have a gas dehydration unit for the purpose of removal of associated water and dew-pointing it to pipeline sales gas specification. In the United States, sales gas transmission pipelines have water specifications of a maximum of 7 lb per MMSCF, while in Canada it is <= 5 lb per MMSCF Dehydration using TEG as a liquid desiccant is the most common and economical means of bulk water removal from natural gas. Thousands of TEG dehydration units have been installed worldwide since the first TEG dehydration unit built in the early part of the 20th century. Natural Gas TEG dehydration is a mature technology and standard configurations are available for a wide variety of wet natural gas compositions.

The purpose of this blog entry is to provide a brief troubleshooting checklist which can prove useful to TEG dehydration unit field operators and maintenance personnel. While the attached checklist does not proclaim to be the one-and-all troubleshooting guide for TEG dehydration units it certainly gives an overview of the most commonly encountered problems and the general checks and actions required to eliminate them. The checklist does not substitute the experience of seasoned operators and maintenance personnel who have had years of operating and troubleshooting experience of TEG dehydration units.

The checklist is attached as both as a picture file and an excel workbook.

I hope to receive some comments and observations from the learned members of "Cheresources".

Regards,
Ankur.

Attached Thumbnails

Attached Files

TEG_Dehydration_Unit_Troubleshooting_Checklist.xlsx (11.92KB)
downloads: 1072

Heat Transfer In Buried Liquid Pipelines

Mon, 14 Dec 2015 13:59:00 +0000

Dear All,

Today's blog entry relates to heat transfer from liquids flowing in long-distance buried pipelines. When liquids have to be transported at relatively high temperatures, e.g. viscous crudes/products, temperatures and cooling rates can be determined using the following formulae:

where:
T_L = average temperature in cross-section of pipeline at distance L, K
T_s = soil temperature at pipeline depth, K
T_inlet = temperature at pipeline inlet (L = 0), K
U_eff = Effective Heat Transfer Coefficient, W/m²-K
d = pipe internal diameter, m
L = pipeline distance at which T_L is to be measured, m
m = mass flow rate of liquid, kg/s
C_p =Specific Heat Capacity of pipeline contents, J/kg-K

Calculation of U_eff

U_convective
For turbulent flow (Re > 4000)

For Laminar flow (Re < 2300)

where:
U_convective = Convective heat transfer coefficient, W/m²-K
λ_liq = liquid thermal conductivity, W/m-K (for crude oil, λ_liq = 0.13 W/m-K)
u = average liquid velocity, m/s
ρ = liquid density, kg/m³
d = pipe internal diameter, m
C_p = Specific Heat Capacity of pipeline contents, J/kg-K
ν = liquid kinematic viscosity, m²/s
g = acceleration due to gravity = 9.81 m/s²
α = thermal expansion coefficient, 1/K (for crude oil, α = 8*10^-4 K-1)
µ_b = bulk liquid viscosity, Pa.s
μ_w = liquid viscosity at wall temperature, Pa.s
T_b = bulk temperature, K
T_w = inside wall temperature, K

Notes:
1. The inside wall temp. (T_w) is generally assumed 2-3 degrees lower than the bulk temp. (T_b) in laminar flow
2. The bulk liquid viscosity (μ_b) and the liquid viscosity at wall temperature (μ_w) can be assumed same for turbulent flow thus reducing the term (μ_b / μ_w) to 1

Specific Heat Capacity of Crude Oil

where:
C_p = Specific heat capacity at temperature T, J/kg-K
T = Temperature, ⁰C
ρ = crude oil density at temperature T, kg/m³

U_wall

where:
U_wall = steel wall heat transfer coefficient, W/m²-K
λ_steel = thermal conductivity of steel, W/m-K (for Carbon Steel, λsteel = 52 W/m-K)
d_o = outside diameter of the steel pipe, m

d = pipe internal diameter, m
t_wall = wall thickness, m

U_coating(Note: External Coating could be a polyolefin liner or insulation)

where:
U_coating = coating heat transfer coefficient, W/m²-K
d_o = outside diameter of the steel pipe, m
λ_coating = thermal conductivity of coating or insulation, W/m-K
t_coating = coating thickness, m

U_environment (For Buried pipes the environment is soil)

where:
U_environment = heat transfer coefficient to the environment (soil), W/m²-K
λ_soil = thermal conductivity of soil, W/m-K
z = burial depth of the pipe up to the pipe axis, m

I have developed a spreadsheet for temperature drop in a buried crude oil pipeline based on the above methodology. For a set of conditions that I have used and for a pipeline inlet crude oil temperature of 55⁰C, the temperature drops to 48.5⁰C at a distance of 50 km for a 16 inch NPS pipeline based on the above method. The pipeline is also externally coated.

That is all for today's blog entry. Hope to have comments from the knowledgeable members of the forum.

Regards,
Ankur.

Attached Thumbnails

Deoiling Hydrocyclones And Their Performance Prediction

Thu, 12 Nov 2015 19:58:00 +0000

Dear All,

Today's blog entry relates to liquid-liquid hydrocyclones used in produced water treatment in the oil & gas industry.

The static hydrocyclone has rapidly been accepted as a compact and efficient means of removing dispersed hydrocarbons from water. The basic design of the static hydrocyclone is illustrated schematically in the figure below:

The water containing the dispersed hydrocarbons enters the hydrocyclone through a tangential inlet atthe top of the swirl chamber. As the liquids swirl along the hydrocyclone, the centrifugal forces genarated promote the separation of the hydrocarbon and water phases, with the hydrocarbon phase forming a thin core at the centre of the hydrocyclone.

By maintaining a suitable pressure ratio between the clean water outlet stream and the reject oil outlet stream, the geometry of the hydrocyclone will result in the thin hydrocarbon core flowing in a reverse direction, exiting from the top of the swirl chamber. The clean water exits from the tail section of the hydrocyclone.

The following definitions are commonly used for hydrocyclones:
Feed: The oil water stream entering the hydrocyclone
Underflow: The clean water stream exiting from the tail of the hydrocyclone
Reject Stream or Overflow: The concentrated hydrocarbon stream exiting from the head of the hydrocyclone through the reject port
Reject Ratio ®: The ratio of the reject and feed stream volumetric flow rates (R = Q_reject / Q_feed)

Performance Prediction:
Cut Size Diameter:
The cut size diameter is a droplet diameter which can be used to characterize the separation performance of a hydrocyclone.

A common cut size diameter definition is the 50% cut size d₅₀.This is defined as the hydrocarbon droplet diameter which has a 50% probability of leaving the hydrocyclone in the reject stream and a 50% probability of leaving the hydrocyclone in the water outlet stream i.e.if

n_{d, reject} / n_d,out = 1 then d = d₅₀

Another commonly used cut size is d75, representing the hydrocarbon droplet size with a 75% chance of being removed from the water stream.

Cut Size Diameter Correlations:
Several correlations are available for the prediction of the droplet cut size. The following correlations are presented:

Bradley Equation:

d₅₀ = 0.053*D_cycl*(ρ_c / (Re_i*Δρ))^0.5

where:
d₅₀ = droplet size diameter as defined above, m
D_cycl = Diameter of the cylindrical hydrocyclone chamber, m
ρ_c = Density of the continuous phase (water), kg/m³
Δρ = Density diff. between the continuous phase (water) (ρ_c) and the dispersed phase (oil) (ρ_oil), kg/m³
Note: The min. recommended density difference for a hydrocyclone is 50 kg/m³ @operating temperature

Re_i = D_i*V_i*ρ_c / μ_c

where:
Re_i = Reynolds number at the hydrocyclone inlet, dimensionless
D_i = inlet diameter at hydrocyclone inlet, m
V_i = Fluid velocity at inlet, m/s
μ_c = viscosity of the continuous phase (water), kg/m-s

Rietema Equation:

d₅₀ = 0.51*D_cycl* ^0.5

where:
d₅₀ = droplet size diameter as defined above, m
D_cycl = Diameter of the cylindrical hydrocyclone chamber, m
ρ_c = Density of the continuous phase (water), kg/m³
Δρ = Density diff. between the continuous phase (water) (ρ_c) and the dispersed phase (oil) (ρ_oil), kg/m³
Note: The min. recommended density difference for a hydrocyclone is 50 kg/m³ @operating temperature

Re_i = D_i*V_i*ρ_c / μ_c

where:
Re_i = Reynolds number at the hydrocyclone inlet, dimensionless
D_i = inlet diameter at hydrocyclone inlet, m
V_i = Fluid velocity at inlet, m/s
μ_c = viscosity of the continuous phase (water), kg/m-s

Coleman-Thew Empirical Model:

d₇₅ = (Hy₇₅*D_cycl ^3*μ_c / (Q*Δρ))^0.5
where:
d₇₅ = droplet size diameter as defined above, m
D_cycl = Diameter of the cylindrical hydrocyclone chamber, m
ρ_c = Density of the continuous phase (water), kg/m³
μ_c = viscosity of the continuous phase (water), kg/m-s
Q = Volumetric flow rate at inlet of hydrocyclone, m³/h
Δρ = Density diff. between the continuous phase (water) (ρ_c) and the dispersed phase (oil) (ρ_oil), kg/m³
Note: The min. recommended density difference for a hydrocyclone is 50 kg/m³ @operating temperature

Hy₇₅ = c₁*(Re_D)^c₂
where:
Re_D = 4*Q*ρ_c / (π*D_cycl*μ_c)
Hy₇₅ = hydrocyclone number, dimensionless
Re_D = Hydrocyclone reynolds number, dimensionless
ρ_c = Density of the continuous phase (water), kg/m³
c₁ & c₂ = Empirically derived constants based on the hydrocyclone model

Some constants are tabulated below for some vendor models

This is all for today. Questions and comments from the members of "Cheresources" are welcome.

Regards,
Ankur.

Attached Thumbnails

Generalized Equation For Crude Oil Working Storage In Refinery Based On Various Factors

Sun, 11 Oct 2015 13:18:00 +0000

Dear All,

Refinery storage of raw materials (mainly crude oil) and finished products is a complex study and involves concepts of linear programming and Monte Carlo simulation. Refer the links below to know about these in context of raw material and product storage:

http://egon.cheme.cmu.edu/Papers/TankFarmTerrazasGrossmannWassick.pdf

http://informs-sim.org/wsc09papers/213.pdf

http://www.palisade.com/risk/monte_carlo_simulation.asp

http://citeseerx.ist.psu.edu/viewdoc/download?doi=10.1.1.132.1037&rep=rep1&type=pdf

However, crude oil storage in refinery can be put in the form of a simplified equation considering various factors such as: Maximum Cargo Capacity (crude parcel size), time duration, Crude Distillation Unit throughput and strategic stockpiling.

Today's blog entry provides such a simplified equation encompassing the above factors to get a preliminary estimate of crude oil storage in a refinery. The equation can be written as follows:

V_c = V_t + (D_e + D_d + D_u + D_s + D_l)*CDU_t + S

where:
V_c = Total Working Capacity of Crude of the Refinery, m³ (bbl)
V_t = Maximum Crude Oil Tanker Cargo or Parcel Size, m³ (bbl)
Note: For large refineries, VLCC tankers with volumetric capacities of 318,000 m³ (2,000,000 bbl) are often the preferred method of receiving crude oil due to the scale of economies.
D_e = Days of early arrival (normally a term that is considered zero in the above equation)
D_d = Days of Delayed Arrival (typically 3 days are considered)
D_u = Days for tanker unloading (typically 1 day)
D_s = Days for settling (typically 1 day)
D_l = Days for leaving, dewatering and crude oil analysis (typically 1 day)
CDU_t = CDU daily throughput i.e. plant capacity, m³/SD (BPSD)
S = Strategic Stockpiling Required (if any), m^₃ (bbl)

Crude oil tankers are planned to enter the port at regular intervals, but may be delayed for bad weather conditions or other reasons. Thus, if tankers are behind schedule, this period must be covered by the crude oil in stock. A maximum of three (3) days should be considered for this period. On the other hand, as the early arrival of a tanker can be controlled by adjusting the cruising speed, it may not be necessary to allow for early arrival in the normal situation.

Individual tank storage capacity and number of storage tanks should be decided on other considerations such as plot plan, pumping capacities to the CDU, limitations of fabrication of large sized storage tanks and HSE studies such as EIA and QRA. A practical maximum tank size could be 130,000 m³ (818,000 bbl) working capacity.

That is all for today's blog entry. Look forward to comments and observations from members of "Cheresources".

Regards,
Ankur.

Distance Between Mist Eliminator Top And Vapor Outlet Nozzle For Vertical Separators

Fri, 04 Sep 2015 09:03:00 +0000

Dear All,

Besides defining the Low-Low (LLL), Low (LL), Normal (NLL), High (HLL) and High-High (HHLL) liquid level based on criteria such as retention time, surge requirements for normal process control and adequate disengagement space during events such as large liquid slugs entering the separator or foaming, it is also important to define the vapor outlet nozzle distance from the top of the mist elimination device.

The objective of providing a minimum distance between the mist elimination device and the vapor outlet nozzle is to prevent channeling by promoting equal velocities across the entire surface area of the mist elimination device.

Refer the sketch below for two separate configurations of mist eliminator and vapor outlet nozzle for vertical gas-liquid separators:

The vertical distance X₄ is the minimum distance required between the mist eliminator top and the outlet nozzle and is given by:

X₄ ≥ (D - d₂) / 2 -------------(1)

Also

h ≥ d₂ ------------------(2)

where:
X₄ = vertical distance from mist eliminator top to vapor outlet nozzle as shown in sketch
D = Outside diameter of vessel
d₂ = Outside diameter of vapor outlet nozzle

Some standards recommend that X₄ should be a minimum of 12 inches even if the calculated value as per equation-1 is less than this value.

That is all for today folks and I would be happy to have comments and experiences from the "Cheresources" community members.

Regards,
Ankur.

Attached Thumbnails

Pressure Drop In Fixed (Single Phase Tubular Packed) Bed Reactors

Wed, 29 Jul 2015 12:33:00 +0000

Dear All,

Today's blog entry relates to pressure drop in fixed bed reactors.

Let us examine the components that contribute to the calculated pressure drop in fixed bed reactors. The design pressure drop should consider a safety margin on the calculated pressure drop.

Bed Pressure Drop: For both gases and liquids this should be calculated based on the "Ergun" equation given as:

(ΔP/L) = K*Re*(150 + 1.75*Re)*((1 – ε) / ε)^3*(µ^2 / (ρ*D_p^3))

where:
(ΔP/L) = pressure drop kPa/m (psi/ft) of bed
Re = Reynolds number, dimensionless = W*D_p/(µ*(1 - ε))
ε = Void fraction of bed, dimensionless

µ = viscosity of gas or liquid at conditions, Kg/m.s (lb/ft.hr)
K = Dimensional constant = 0.001 (Metric) = 1.665E-11 (USC)
ρ = gas or liquid density at conditions, kg/m³ (lb/ft³)
D_p= Equivalent particle diameter, m (ft)
W = mass flux of gas or liquid, kg/s.m² (lb/hr.ft²) based on reactor cross-sectional area

The equivalent particle diameter to be used in pressure drop calculations for cylindrical shapes is given by the following:

D_p = 3*D*L / (2*L + D)

where:
D_p = Equivalent particle diameter, m (ft)
D = Actual particle diameter, m (ft)
L = Average particle length, m (ft)

The "Ergun" constants of 150 and 1.75 have been given by fitting pressure drop data from a wide variety of particle systems. Use of the equation and these constants for a given unique shaped particle can sometimes result in significant error in the pressure drop prediction. In addition, error is often introduced into the pressure drop calculation through the estimate of bed voidage and the equivalent particle diameter. These errors are caused due to uncertainty in the particle density and particle dimensions.

Sock Loading: Prior to the 1970s, the standard method for loading catalyst in a fixed bed reactor was sock loading. In sock loading, a canvas tube conveys the catalyst from the reactor inlet manway to the bottom of the reactor catalyst bed. The sock is attached to a loading hopper or funnel at the reactor inlet, which discharges the catalyst through the sock upon the bed surface in a manner which prevents the individual particle (specifically cylinders) from finding a stable, horizontal rest position. The cylinders stack in various horizontal and vertical positions.

The positioning of catalyst cylinders in random orientations encourages bridging of cylinders and void spaces between cylinders. During reactor operations, these bridges and void spaces tend to collapse. Bed density then increases as the bed depth shrinks.

Dense Loading: Since 1970, refiners, catalyst manufacturers, and catalyst-loading contractors have developed dense-loading devices that dramatically reduce void spaces and bridging. Dense loading can increase catalyst bed densities by as much as 17%.

Moreover, unlike sock loading, dense loading does not require personnel inside the reactor to distribute the catalyst evenly from the sock. Workers inside the reactor require breathing air and weight distribution shoes to prevent crushing of the catalyst underneath their weight.

Dense loading is accomplished by introducing the catalyst particles (specifically cylinders) into the reactor in a manner that allows each cylinder to fall freely to the catalyst surface. Individual cylinders separately assume a horizontal rest position before being impinged by other cylinders. Under this regime, cylinders tend to pack horizontally, minimizing the possibility of bridging or creating void spaces.

Bed Lifting Pressure Drop: For the less frequently encountered upflow operation in a reactor, the calculated pressure drop from the above equation must be kept below the theoretical bed lifting pressure drop, preferably less than 50% and in no case more than 75%. The bed lifting theoretical pressure drop is given by:

(ΔP/L)_BL = K*(ρ_p – ρ)*(1 – ε)

where:
(ΔP/L)_BL = Theoretical Bed Lifting Pressure Drop of bed, kPa/m (psi/ft)
ρ_p = particle density of catalyst, kg/m³ (lb/ft³)
ρ = Density of gas or liquid at conditions, kg/m³ (lb/ft³)
ε = Void fraction of bed, dimensionless
K = Dimensional constant = 0.0098 (Metric) = 0.00694 (USC)

Inlet Nozzle Pressure Drop: Pressure drop through the inlet nozzle is the sum of the following terms represented in equation form:

ΔP_e = (K*ρ*(U_L – U_I)^2) / 2
ΔP_i = (1.3*K*ρ*U_I^2) / 2
ΔP_b = (0.5*K*ρ*U_L^2) / 2
where:
ΔP_e = pressure loss due to sudden expansion from line to expanded section of inlet distributor, kPa (psi)
ΔP_i = pressure drop due to impingement of the gas in the bottom plate or dish of the inlet distributor, kPa (psi)
ΔP_b = pressure drop through the slots of the inlet distributor, kPa (psi)
ρ = Density of gas or liquid at conditions, kg/m³ (lb/ft³)
U_I = velocity in expanded section of inlet distributor, m/s (ft/sec)
U_L = velocity in external piping, m/s (ft/sec)
K = Dimensional constant = 0.001 (Metric) = 2.156E-4 (USC)

Collector and Outlet Nozzle Pressure Drop: Pressure drop through the collector and outlet nozzle is the sum of the following terms represented in equation form:

ΔP_s = (2.8*K*ρ*U_S^2) / 2
ΔP_c = (0.5*K*ρ*U_L^2) / 2
where:
ΔP_s = Pressure drop through holes and slots of collectors, kPa (psi)
ΔP_c = Pressure drop due to sudden contraction at entrance to the outlet nozzle, kPa (psi)
U_S = velocity through holes and slots, m/s (ft/sec)
K = Dimensional constant = 0.001 (Metric) = 2.156E-4 (USC)

This is all for today's blog entry and I look forward to comments from the reader's of my blog.

Regards,
Ankur.

Attached Thumbnails

Zeolite Based Molecular Sieve Adsorbents In The Chemical Process Industry

Sat, 20 Jun 2015 04:52:00 +0000

Dear All,

Today's blog entry gives a brief overview of zeolite-based molecular sieves for physical adorption operations in the chemical process industry.

Since the 1960s, molecular sieve adsorbents have become firmly established as a means of performing difficult separations, including gases from gases, liquids from liquids and solutes from solutions. They are supplied as pellets, granules, or beads and occassionally as powders. The adsorbent may be used once and discarded or, more commonly may be regenerated and used for many cycles over an extended time period. They are generally stored in cylindrical vessels through which the stream to be treated is passed. For regeneration, two or more beds are usually employed with suitable valving, in order to obtain a continuous process. As a unit operation adsorption is unique in several respects. In some cases, one separation is equivalent to hundreds of mass transfer units. In others, the adsorbent allows the selective removal of one component from a mixture, based on molecular size differences, which would be extremely unlikely to be performed by any other means. In addition, contaminants can be removed from fluid streams to attain virtually undetectable impurity concentration. Adsorbents may be used in applications for adsorbing a few grams to several tons.

Before focusing on zeolite adsorbents, it is important to understand the two basic categories of adsorbents i.e. Amorphous and Crystalline.

Amorphous Adsorbents: The amorphous adsorbents (silica gel, activated alumina, and activated carbon) typically have specific areas in the 200-1000 m²/g range, but for some activated carbons much higher values have been achieved (~1500 m²/g). The difficulty is that these very high area carbons tend to lack physical strength and this limits their usefulness in many practical applications. The high area materials also contain a large proportion of very small pores, which renders them unsuitable for applications involving adsorption of large molecules. The distinction between gas carbons, used for adsorption of low molecular weight permanent gases, and liquid carbons, which are used for adsorption of larger molecules as color bodies from the liquid phase, is thus primarily a matter of pore size.

In a typical amorphous adsorbent the distribution of pore size may be very wide, spanning the range from a few nanometers (nm) to perhaps one micrometer. Since different phenomena dominate the adsorptive behavior in different pore size ranges, IUPAC has suggested the following classification:
micropores, <2 nm diameter
mesopores, 2-50 nm diameter
macropores, >50 nm diameter

The division is somewhat arbitrary since it is really the pore size relative to the size of the sorbate molecule rather than the absolute pore size that governs the behavior. Nevertheless, the general concept is useful. In micropores (pores which are only slightly larger than the sorbate molecule) the molecule never escapes form the force field of the pore wall, even when in the center of the pore. Such pores generally make a dominant contribution to the adsorptive capacity for molecules small enough to penetrate. Transport within these pores can be severely limited by steric effects, leading to molecular sieve behavior.

The mesopores make some contribution to the adsorptive capacity, but their main role is as conduits to provide access to the smaller micropores. The macropores make very little contribution to the adsorptive capacity, but they commonly provide a major contribution to the kinetics. Their role is thus analogous to that of a super highway, allowing the adsorbate molecules to diffuse far into a particle with a minimum of diffusional resistance.

Crystalline Adsorbents: Crystalline adsorbents are the zeolites and zeolite analogues such as silicalite and the microporous aluminum phosphates. where the dimensions of the micropores are determined by the crystal framework and there is therefore virtually no distribution of pore size. However, a degree of control can sometimes be exerted by ion exchange, since, in some zeolites, the exchangeable cations occupy sites within the structure which partially obstruct the pores. The crystals of these materials are quite small (1-5 microns) and they are aggregated by a suitable binder (generally a clay) and formed into macroporous particles having dimensions large enough to pack directly into an adsorber vessel. Such materials therefore have a well-defined bimodal pore size distribution with the intracrystalline micropores (a few tenths of a nanometer) linked together through a network of macropores having a diameter of the same order as the crystal size (~1 micron).

Here it is important to introduce "Desiccants" as a class of adsorbents and the following describes them.

Desiccants: A solid desiccant is simply an adsorbent which has a high affinity and capacity for adsorption of moisture so that it can be used for selective adsorption of moisture from a gas (or liquid) stream. The main requirements for an efficient desiccant are therefore a highly polar surface and a high specific area (small pores). The most widely used desiccants are silica gel, activated alumina and the aluminum-rich zeolites (4A or 13X).

The zeolites have high affinity and high capacity at low partial pressures. This makes them useful desiccants where a very low humidity or dew point is required. The micropores of 3A zeolite are small enough to exclude most molecules other than water. It is therefore useful for drying reactive gases. The major disadvantage of zeolite desiccants is that a high temperature is required for regeneration (>300 deg C), which makes their use uneconomic when only a moderately low dew point is required.

Temperature Swing Adsorption (TSA) Process: A temperature-swing or thermal-swing adsorption (TSA) cycle is one in which desorption takes place at a much higher temperature than adsorption. The principal application is for separation in which contaminants are present at low concentration, i.e for purification. The TSA cycles are characterized by low residual loadings and high operating loadings. These high adsorption capacities for low concentrations means that cycle times are long, hours to days to months, for reasonably sized beds. This long cycle time is fortunate, because packed beds of adsorbent respond slowly to changes in gas temperature. A purge and / or vacuum removes the thermally desorbed components from the bed, and cooling returns the bed to adsorption condition. Systems in which species are strongly adsorbed are especially suited to TSA. Such applications include "Drying", "Sweetening", "CO2 removal", and "pollution control".

A detailed description of the TSA process would require another blog entry and I end it with the description provided above.

Applications for Adsorbents:

Drying: The single most common gas phase application for TSA is drying. The natural gas, chemical, and cryogenic industries all use zeolites, silica gel, and activated alumina to dry streams. Adsorbents are even found in mufflers.

Zeolites, activated alumina, and silica gel have all been used for drying of pipeline natural gas. Alumina and silica gel have the advantage of having higher equilibrium capacity and of being more easliy regenerated using waste level heat. However, the much lower dew point and longer life attainable with 4A makes zeolites the predominant adsorbent. Special acid-resistant zeolites are used for natural gas containing large amounts of acid gases, such as H₂S and CO₂.

The low dewpoint that can be achieved with zeolites is especially important when drying feed streams to cryogenic processes to prevent freeze-up at process temperatures. Natural gas is dried before liquefaction to liquefied natural gas (LNG), both in peak demand and in base load facilities. Zeolites have largely replaced silica gel and activated alumina in drying natural gas for ethane recovery utilizing the cryogenic expander process, and for helium recovery. The air to be cryogenically distilled into N₂, O₂ and argon must be purified of both water and CO₂. This purification is accomplished with 13X zeolites.

The 4A zeolite, silica gel, and activated alumina all find applications in drying synthesis gas, inert gas, hydrocracker gas, rare gases, reformer recycle H₂. Cracked gas before low-temperature distillation for olefin production is a reactive stream. The 3A or pore-closed 4A zeolite selectively adsorbs water but excludes the hydrocarbons, thus preventing coking. It also prevents co-adsorption of hydrocarbons which would otherwise be lost during desorption with the water. Small pore zeolites are also applied to the drying of ethylene, propylene, and acetylene as they are drawn from storage. When industrial gases containing Cl₂, SO₂ and HCl are dried, acid-resistant zeolites are used.

A relatively new application for zeolites is the prevention of corrosion in automotive mufflers.

Sweetening: Another significant purification application area for adsorption is sweetening. H₂S, mercaptans, organic sulfides and disulfides, and COS need to be removed to prevent corrosion and catalyst poisoning. They are found in H₂, natural gas, de-ethanizer overhead, and bio-gas. Often adsorption is attractive because it dries the stream as it sweetens.

In the sweetening of wellhead natural gas to prevent pipeline corrosion, 4A zeolites allow sulfur compounds removal without CO₂ removal (to reduce shrinkage), or the removal of both to upgrade low thermal content gas. When minimizing the formation of COS during desulfurization is desirable, calcium-exchanged zeolites are commonly used because they are less catalytically active for the reaction of CO₂ with H₂S to form COS and water. Natural gas for steam-methane reforming in ammonia production must be sweetened to protect the sulfur-sensitive, low temperature shift catalyst. Zeolites are better than activated carbon because mercaptans, COS, and organic sulfides are also removed. Many refinery H₂ streams require H₂S and water removal by 4A and 5A zeolites to prevent poisoning of catalysts such as those in catalytic reformers.

Other Separations: Other TSA applications range from CO₂ removal to hydrocarbon separations, and include removal of air pollutants and odors, and purification of streams containing HCl and boron compounds. Because of their high selectivity for CO₂ and their ability to dry concurrently, 4A, 5A, and 13X zeolites are the predominant adsorbents for CO₂ removal by temperature-swing process. The air fed to an air separation plant must be H₂O and CO₂ free to prevent fouling of heat exchangers at cryogenic temperatures; 13X is typically used here. Another application for 4A-type zeolite is for CO₂ removal from base load and peak-shaving natural gas liquefaction facilities.

Zeolites with high acid resistance, such as mordenite and clinoptilolite, have proven to be effective adsorbents for dry SO₂ removal from sulfuric acid tail gas, and special zeolite adsorbents have been incorporated into the UOP PURASIV S process for this application.

Zeolites have also proven applicable for removal of nitrogen oxides (NO_x) from wet nitric acid plant tail gas by the UOP PURASIV N process.

Mordenite and Clinoptilolite zeolites are used to remove HCl from Cl₂, chlorinated hydrocarbons, and reformer gas streams.

Refer the link below for details on application of zeolite molecular sieves in various areas of the chemical process industry:

Which_Zeolite_To_Use

Additionally refer the attached table to get an overview of the commonly used zeolite molecular sieves in the chemical process industry:

While making an attempt to make this blog entry as short as possible, I feel the information provided is a bare minimum to get an overview of how zeolite based molecular sieves are useful in adsorption processes. I look forward to comments from the esteemed members of the "Cheresources" community on this blog entry.

Regards,
Ankur.

Attached Thumbnails

Skid Mounted Modular Mini Crude Refining Units

Thu, 28 May 2015 06:15:00 +0000

Dear All,

In remote crude oil drilling locations or where some value addition on the crude is required such as production of low cetane diesel (generally less than 40) and straight run naphtha with low first investment costs, skid mounted modular mini refining crude units are the answer.

Let us examine the advantages of a skid mounted modular unit:

1. Lower construction and installation costs

2. Ideal choice for remote locations

3. Less manpower requirement for installation, operation and troubleshooting of the unit

4. Very short installation, commissioning and operation times. A skid mounted modular crude distillation unit with a capacity of 300 to 12,000 BPD could be made operational within weeks from it's arrival at the designated location.

5. Since the units are skid mounted they are easily transportable and can be moved from one location to another in a short duration of time. This feature allows the units to be procured on short or long term lease basis thereby further reducing cost for the operator and increasing profitability by moving the unit to the production source as and when desired.

6. Higher quality workmanship can be achieved in view of the fact that these units are assembled in sophisticated workshops utilizing high quality machinery and considering that these may have to face the rigours of frequent transportation depending on the operator's philosophy of operation.

7. Construction as modules allows transport by sea in standard ocean containers.

I am providing a list of some established names who design, build, install and commission such units. Links are provided for some of the leading suppliers of the skid mounted modular crude mini crude refining units for the benefit of the readers.

http://www.skmind.com/SKM_Diesel_Refinery_Brochure_.pdf

http://mgiprocess.com/modular-refineries.htm

http://chemexmodular.com/Modular_Refineries

http://www.petrostarpetroleum.com/downloads/Farrow-Refinery-Summary.pdf

As a cautionary advice the credentials of the suppliers of these units needs to be verified as a whole in terms of engineering, workshop fabrication of skid mounted modules, skilled pre-commissioning and commissioning personnel and after-sales service. It would be unwise for a prospective client to try to procure such units from a source where there are multiple agencies involved for engineering, fabrication and pre-commissioning / commissioning, due to the possibility of encountering fly-by-night operators who are out to make a fast buck at the client's peril. The most prudent way for a prospective client would be to check out the performance track record of the supplier in terms of quantity and quality of units supplied in the past and proper testimonials from satisfied clients.

Comments from the readers of my blog are welcome.

Regards,
Ankur.

Overall Heat Transfer Coefficient For Tank Heaters (Steam) Heating Hfo / Asphalt

Wed, 29 Apr 2015 18:51:00 +0000

Dear All,

Today's blog entry provides an equation for Overall Heat Transfer Coefficient for storage tanks handling Heavy Fuel Oil (HFO) / Asphalt and provided with tank heating coils using steam as a heating medium.

If the heat loss rate Q (W or Btu/hr), from the tank is known, then the steam coil surface area can be calculated and thus the coil geometry fixed. The general equation for sizing heat exchange equipment can be utilized for this:

Q = U_o*A*ΔT

or

A = Q /( U_o*ΔT)

where:
A = Steam Coil surface area, m² or ft²
Q = Heat Loss Rate from tank, W or Btu/hr
U_o = Overall Heat Transfer Coefficient, W / m².⁰C, Btu / hr.ft².⁰F
ΔT = Temperature difference between steam supply temperature and bulk process fluid temperature in tank, ⁰C, ⁰F

For heat loss rate from storage tanks, refer this very well known post on "Cheresources":

http://www.cheresources.com/invision/topic/9066-storage-tank-heat-loss-calculation-based-on-kumana-and-kothari-article/?hl=kumana

Now for the equation for the equation for the Overall Heat Transfer Equation, U_o

United States Customary Units:

U_o = 30*(T_c - T_f)^0.14 / (µ_f)^0.4

Metric Units:

U_o = 11.7*(T_c - T_f)^0.14 / (µ_f)^0.4

where:

T_c = Temperature of the steam, ⁰F, ⁰C
T_f = Bulk Temperature of the process fluid, ⁰F, ⁰C
µ_f = The absolute viscosity of the process fluid at the average film temperature ((T_f + T_i) / 2), cP, Pa.s

where:
T_i = Temperature of the coil surface in contact with the fluid, ⁰F, ⁰C

Generally T_i is 10-15⁰C more than the bulk process fluid temperature

Once the steam coil surface area A is calculated, the length of the coil can be calculated after fixing a NPS or OD of the steam coil pipe.

That is it for today's blog entry. I look forward to comments from the members of "Cheresources" on this blog entry.

Regards,
Ankur.

Safety Integrity Level (Sil) Definition And Brief Explanation

Sun, 29 Mar 2015 08:10:00 +0000

Dear All,

Most young process engineers have heard the term SIL but not got involved in what SIL is all about. In fact, there is a misconception among many younger process engineers that SIL is solely related to the advanced "Control & Automation" part of a process plant and process engineers need not get involved in a detailed SIL study or SIL review except for providing process data for the instrumentation under SIL study.

This is far from the truth. Process engineers need to be an integral part of any SIL review or SIL study because the basis or starting point of any SIL study is proper evaluation and finalization of the "Basic Process Control System" (BPCS) based on process studies / reviews such as "Design Review" and "Hazard & Operability Studies" (HAZOP).

SIL studies and SIL allocation for any process plant is a logical step ahead of the BPCS for safe and reliable plant operation.

Now, that I have explained how it is important for process engineers to be part of a SIL study exercise, let us get to the definitions of various terms and explanation of the methodology of SIL.

Some basic terms:

1. Probability of Failure on Demand (PFD): It It is a measure of safety system performance in terms of the Probability of Failure on Demand (PFD). It is expressed as a negative exponential of 10, for example, 10^-5 .

2. Risk Reduction Factor: This is the inverse of of the POF and provides the reduction in risk by implementation of a SIL level to any critical safety-related instrumentation.

3. Safety-Instrumented Systems: It is a process plant instrument system which is designed to prevent or mitigate hazardous events by taking a process to a safe state when predetermined conditions are violated. Other common terms for SIS are safety interlock systems, emergency shutdown systems (ESD), and safety shutdown systems (SSD).

SIL evaluation is done for Safety-Instrumented Systems (SIS). Each SIS has one or more Safety Instrumented Functions (SIF). To perform its function, a SIF loop has a combination of logic solver(s), sensor(s), and final element(s). Every SIF within a SIS will have a SIL level. These SIL levels may be the same, or may differ, depending on the process. It is a common misconception that an entire system must have the same SIL level for each safety function.

SIL Levels (as per IEC 61508):

There are four discrete integrity levels associated with SIL: SIL 1, SIL 2, SIL 3, and SIL 4. The higher the SIL level, the higher the associated safety level, and the lower probability that a system will fail to perform properly. As the SIL level increases, typically the installation and maintenance costs and complexity of the system also increase. Specifically for the process industries, SIL 4 systems are so complex and costly that they are not economically beneficial to implement. Additionally, if a process includes so much risk that a SIL 4 system is required to bring it to a safe state, then there is a fundamental problem in the process design that needs to be addressed by a process change or other non-instrumented method.

The identification of risk tolerance is subjective and site-specific. The owner / operator must determine the acceptable level of risk to personnel and capital assets based on company philosophy, insurance requirements, budgets, and a variety of other factors. A risk level that one owner determines is tolerable may be unacceptable to another owner. Many well-known operating companies have their in-house guidelines / standards for assigning SIL levels for specific SIS and for a given type of process plant / unit.

As part of the engineering design cycle, SIL or SIS studies are conducted in a manner similar to other studies such as Design Review & HAZOP. The SIL or SIS study basic prerequisite is, availability of P&IDs and Plant Operation / Control / Safeguarding philosophy, which have been updated to incorporate all process design and HAZOP review comments.The SIL or SIS study is jointly driven by the process engineering and Control & Automation (Instrumentation) engineering group. A HSE engineer also plays an important part in this study / review. Based on the recommendations and report of the SIS study, SIL levels are assigned to various SIS systems in the process plant and instrument system architecture for these SIS systems defined.

The finer details of how the system architecture for the SIS needs to be constructed is handled by the Instrumentation Engineer and is beyond the scope of this blog article.

A basic explanation on SIL is also provided on the links provided below:

http://en.wikipedia.org/wiki/Safety_Integrity_Level

http://www.enggcyclopedia.com/2011/10/safety-integrity-levels-sil/

A google search will provide many more articles related to SIS/SIL.

Hope this blog article helps process engineers to understand the basics of SIL and SIS. Looking forward to comments from the reader's of my blog.

Regards,
Ankur.

Attached Thumbnails

Nit Warangal, Department Of Chemical Engineering Golden Jubilee Celebrations

Sun, 15 Mar 2015 15:25:00 +0000

Dear All,

It is heartening to note that the contribution of "Cheresources" to the global chemical engineering community is being recognized all over including academic institutions.

National Institute of Technology (NIT), Warangal, is one of the premier engineering institutes of India providing bachelors, masters and doctoral degrees in the field of engineeering, technology and management.

NIT Warangal is a federally funded institute by the Government of India and ranks among the top fifteen technological institutes of India. For more details on NIT Warangal, refer the following link:

http://en.wikipedia.org/wiki/National_Institute_of_Technology,_Warangal

It recently (March 13-14) concluded a national level student's conference named "ChemFlair" conducted by its Department of Chemical Engineering on the occasion of completing 50 years from its inception.

It was a singular honor for me, to be invited as a Chief Guest on this occasion as a practicing chemical engineer. The chemical engineering department of NIT-W has a very distinguished faculty with most of them holding Doctoral degrees from premier engineeeing institutes of India such as IITs and from top notch universities in the United States.

Since it was a undergraduate student's conference, chemical engineering undergraduate student's were invited from all over the country to participate in discussions, give presentations, participate in Q&A sessions and various technical quizzes.

I made a presentation titled "Phases of Process Design Engineering" from the perspective of an engineering consultant and contractor. The idea was to give the undergrad's a bird's eye view on the practical aspects of engineering when working in a business environment of engineering, procurement and construction of chemical process plants.

The interaction with the faculty and the student's made me realize that at the undergraduate level student's are quite unaware of practical business scenarios involving engineering of a chemical process plant and my presentation helped them in clearing quite a few gaps in their understanding. One another aspect that I could make them understand about engineering design was implementing a safe and environment friendly process design. While most of them were aware of Process Flow Diagrams, Process Modeling and Simulation, Heat and Mass Balance, P&IDs, Process Datasheets, Equipment Sizing, almost none of them had awareness of safety studies conducted during the engineering design phase, such as HAZID, HAZOP, SIL etc. For many of them, my explanation on the safety aspect of engineering design was a revelation.

During the inaugural address and while being introduced to the particpant's of the conference, I was understandably delighted to see the "Community Index" page of "Cheresources" and "Ankur's Blog" being displayed on the screen. It was good to see that "Cheresources" contribution to our community was being acknowledged by the academia.

While many of the younger generation of internet savvy student's were aware of "Cheresources", they had only gone through the forums on a very superficial basis. After my interaction with them, many of them expressed their desire to increase their involvement with "Cheresources".

To conclude, it was a great honor to be part of the student's conference organized by one of the premier engineering institutes of India and more so to interact with the younger generation of engineers and would-be engineers who will guide the future of chemical engineering in the years to come.

Would love to hear comments from the readers of my blog.

Regards,
Ankur.

Requirement Of Adequate Ventilation In Process Areas

Wed, 18 Feb 2015 17:30:00 +0000

Dear All,

Process Engineers normally are not directly involved in Heating, Ventilation, Airconditioning (HVAC) requirements of a Chemical Process Plant. HVAC engineering is a specialized branch of engineering and HVAC engineers normally involve themselves in system design for a chemical process plant. However, adequate ventilation is a matter of concern for all engineers who consider safety as an integral part of any chemical process plant and it is important to know the hazards related to inadequate ventilation.

NFPA 30 defines adequate ventilation from the perspective of a chemical process plant as follows:

An area is adequately ventilated if it is ventilated at a rate sufficient to maintain the concentration of vapors within the area at or below 25% of the lower flammable limit.

This shall be confirmed by one of the following:
- Calculations based on the anticipated fugitive emissions; or
- Actual vapor concentration sampling under normal operating conditions, conducted at a radius of 5 feet from equipment.

An acceptable alternative is to provide ventilation at a rate of not less than 1 cubic foot per minute per square foot of solid floor area.

Ventilation shall be accomplished by natural or mechanical ventilation, with discharge of exhaust to a safe location, without recirculation of the exhaust air.

Completely open (from all 4 sides) outdoor locations are usually considered adequately ventilated.

Process equipment buildings are buildings that contain hydrocarbon piping and equipment. Some examples are: analyzer buildings, metering stations, pump stations, compressor stations, and separation stations. They should have openings on all sides whenever practical to allow for ventilation. If they must be fully enclosed due to ambient conditions, adequate ventilation needs to be provided. to prevent the accumulation of flammable gases. The building exhaust vents should be located to preclude the accumulation of gases at low points such as pipe trenches, as well as at the ceiling.

Process buildings are frequently ventilated with windows, floor louvers, and roof ventilators.

A minimum of six air changes per hour are recommended for process buildings.

If mechanical equipment provides the required ventilation, safeguards are needed to protect against its undetected failure.

Discharge or exhaust must be to a safe location outside the building.

Ventilation rates can be achieved either by continuous introduction of fresh makeup air into the enclosure, or by recirculation of air in the building. A recirculation system should ensure that the air is monitored continuously using a system that automatically alarms, stops recirculation, and provides full exhaust to the outside in the event that vapor-air mixtures over 25% of the lower flammable limit are detected.

Recirculation should be designed with adequate air movement and direction to minimize “dead” areas where flammable vapors such as heavy hydrocarbons may collect.

If conditions include the risk of a large flammable vapor release occurring in a confined space, and the calculated rate of diluent ventilation is not sufficient to dilute and disperse the released vapor to below the "Lower Flammable Limit" (LFL) within a reasonable time (i.e., four hours), then supplemental emergency ventilation should be provided. Emergency ventilation can be natural ventilation through panels or louvers, or switching recirculation fans to full fresh air make-up or exhaust. The travel direction of ventilated vapor should avoid its reaching an ignition source outside the enclosed space being ventilated.

It is recommended that when combustible gas detectors are installed they provide an alarm at 20% of LFL and are interlocked to rotary process equipments such as pumps and compressors so that they cannot be started or they shut down when the detector detects a value of 60% of LFL.

To conclude, adequate ventilation in a chemical process plant is of utmost importance and process engineers should be aware of good ventilation practices for a safe design.

Hope to have comments from the members of "Cheresources".

Regards,
Ankur.

Don't Be An Engineer At The Expense Of Common Sense

Sun, 04 Jan 2015 10:34:00 +0000

Dear All,

Happy New Year to all of you and I start off with a blog entry which you may find quite interesting.

Being an engineer is justifiably a matter of pride since engineers have made an immeasurable contribution to the betterment of the world we live in.

However, as Homo sapiens, who are the considered on the top of the chain of living beings we must understand that our basic intelligence is the attribute that distinguishes us from other species of the animal kingdom and gives us the hallowed status we have. We have evolved from the stone-age to today's world of gadgetry and gizmos because we have used our intelligence or in other words our common sense more than any other species of the animal kingdom.

While the meaning of common sense can be debated from various perspectives of the human society, I am only going to specifically dwell on common sense from the perspective of a design engineer. Wikipedia has a long article on the various aspects of common sense, which can be read at the link below:

http://en.wikipedia.org/wiki/Common_sense

Coming to specifics, I would like to discuss how common sense and the ability to question and challenge should be used by a design engineer:

1. Certain documents like design basis are prepared at the start of the project. Subsequent engineering design work results from the status quo of the design basis. In other words, the design basis needs to be frozen at the earliest and should not remain dynamic throughout the project unless there are some compelling circumstances. Often, design basis provides listing of future engineering documents that will be prepared during the course of the project. Many of these future documents are identified by an alpha-numeric code and description which is established at the start of the project. A number of such documents identified in the design basis by this code and description are dynamic in nature till the end of the project. An example would be P&IDs for the project. While listing the documents by the alpha-numeric code and the description makes perfect sense in the design basis, providing the revision status of these dynamic document does not. There is absolutely no prudence in revising the design basis because the revision status of the dynamic documents has changed. Engineers must realize that when a listing of a dynamic document is made in a frozen or status quo document, they have the responsibility to refer the latest revision of that document. This is what I mean by common sense.

2. While I do not condone gramatically incorrect or misspelt language, it is important to note that documents such as design and sizing calculations with formulas, equations, sketches do not diminish in stature just because of a few misspelt words or grammatically incorrect language. Anybody rejecting the document solely on these grounds, does need to have lessons in common sense. Obviously, if time is not of essence, which normally is, you can decorate the document with Queen's English and make it as glossy and slick as an advertisement billboard.

3. In a majority of projects either the client or his PMC (Project Management Consultant) checks the engineering deliverables provided by the engineering contractor. Often it is seen that the client or his PMC will put multiple question marks on a particular section of the document, or ask vague and unrelated open-ended questions. I consider this behavior as trying to insult or humiliate the engineering contractor and would treat it with the contempt it deserves. Obviously, people indulging in this kind of antics need extended lessons in etiquette and common sense. It also puts a big question mark in terms of their capabilities as an engineer.

4. E-mail communication betweeen various parties during the project has to be very carefully scrutinized before being sent. Being a client does not give license for using uncouth, foul and patronizing language to the service provider in written communication. As a client, I have had my fair share of grievances with the service provider and many a times I have given them a piece of my mind, however, only on telephone and never by written communication. Remember, the written word has far more grave consequences, specifically when it is bad, than the spoken word. This is one more thing which relates to common sense.

5. Engineering contractors earn revenue and profit based on time management of their skilled personnel. This time management is in terms of their budgeted manhours for the project. From a strictly technical view point, any over-run over the budgeted manhours means a loss of revenue and profits for the engineering contractor. Most clients have little or no concern about the time management of the engineering contractor. They have an expanding wish-list over the tenure of project where they want the engineering contractor to explore all kinds of possibilities and permutations-combinations for the project. If the client is willing to pay for such kind of research and development work, then there is no argument. However, in a majority of cases clients are unwilling to do so and the contractor is burdened with work which is far beyond the scope of work envisaged at the start of the project. While this is undoubtedly unethical on the part of the client, it is also the responsibility of the senior management on the engineering contractor's side to ensure that the client pays for the extra work demanded by him. Businesses do not survive on freebies and charities. This is again common sense for both the client and the service provider.

Well it is time to conclude this somewhat descriptive blog. However, this is not the end of it. I may come out with a Part 2 on the same topic with more common sense related points. Meanwhile, I request the readers of my blog to come out with points related to common sense from an engineering perspective, which they feel can make the world a lot better place for engineers to work.

Regards,
Ankur.

PS: My spouse and a majority of women think that most men have ridiculously low levels of common sense related to home affairs and in managing children. I have no logic to disagree. Do you have a logic to counter this charge? If yes, please share.

Guidance Notes On Buried Piping

Mon, 15 Dec 2014 08:34:00 +0000

Dear All,

Most process engineers should try to understand basic concepts of piping installation along with routine process design calculations related to piping such as pressure drop, velocity and erosion / corrosion limits for various pipe metallurgies.

Today's blog entry provides some guidance on buried or underground pipe installation.

Piping should not be buried or installed underground when it can be reasonably avoided.

Major applications for buried piping are generally cross-country pipelines where security and safety justify burial. Here security implies the difficulty of sabotage and intentional damage to a buried pipeline vis-a-vis an aboveground pipeline. Safety implies the safety of the surroundings (environment / population centers) due to unintentional or accidental damage of the buried pipeline vis-a-vis an aboveground pipeline.

Other major applications for buried piping are firewater piping, critical piping running across the entire plant installation with certain sections aboveground and certain sections underground. Underground sections would generally be routine movement areas including road and rail crossings.

Underground firewater piping has gained popularity in recent years as explained below:

- Above ground carbon steel or lined carbon steel pipes filled with untreated water are a source of corrosion and undermine the integrity of the fire water system and are slowly being replaced with plastic pipes (HDPE / GRE etc.) which are much more corrosion resistant.

-Plastic pipes although having higher corrosion resistance to untreated water have much lower mechanical strength compared to carbon steel pipes when subjected to forces such as impact and vibration.

-To maintain mechanical integrity of the plastic pipe, burying the firewater pipe and running across the installation becomes a logical option.

- Option of having firewater piping with exotic metallurgies (Stainless steels / Cupro-Nickel etc.) is not an economical choice. It is important to note that common austenitic stainless steels are susceptible to chloride stress corrosion cracking in water environment with high chloride content.

Problems Associated with Buried Piping:

- Buried Steel pipes are subjected to external corrosion despite mitigating measures such as external coating and cathodic protection

-Draining, cleaning of buried pipes is difficult compared to aboveground pipe.

-Leak detection and repair of of buried pipe is a difficult and expensive exercise. Although modern buried pipeline leak detection systems are available now, they are very expensive to install.

-Buried pipes are almost always subject to mechanical damage if being excavated or if any excavation work is being carried out in close vicinity.

-Buried pipes carrying hot fluids and subject to thermal expansion can cause pipe deformation and / or partial / total removal of the external protective coating if applied.

Corrosion Protection of Buried Pipes:

All buried steel piping with the possible exception of cast iron piping should be protected from soil corrsion with a suitable exteranl coating.

Following is the list of the most commonly used acceptable coatings and wrappings with approximate pipe surface temperature limitations:

-Fusion Bonded Epoxy (< 93°C)
-Liquid epoxies (< 107°C)
-Extruded Plastic (< 82°C)
-Tape Wraps (< 60°C) (higher temperatures applicable in case of high temperature thermosetting tape)
-Coal Tar Enamels (< 60°C)

When small quantities of buried piping require protection, tape wrap is often selected because it is relatively inexpensive and easy to apply in the field. However, it is often the least reliable coating, with poor performance in water- and oil-saturated soils, and in cyclic temperature service. It requires proper pipe surface preparation and is easily applied improperly.

Protection for coated pipe at weld joints and tie-ins is provided by field-applied fusion bonded epoxy, shrink sleeves of polyethylene, heat-cured liquid epoxy, or tape wrap.

Aside from the type of coating selected, proper application of the coating and maintenance of its integrity are required for proper installation of a protected line. Because success or failure cannot be determined for an extended time after installation, usually years, attention should be paid to:

-Correct surface preparation for the type of coating used
-Application of the coating to the specified consistency and thickness
-Care in handling and laying to avoid coating damage
-Proper cleaning, priming, and field coating of joints and fittings
-Inspection of coating for any damage and proper repair
-Backfilling and compacting to prevent contact with any material that could damage the coating

Cathodic Protection:

Cathodic protection (CP) can be roughly defined as retarding or preventing the corrosion of a metal by imposing an electrical current flowing to the metal through an electrolyte. In the case of buried piping, the pipe is the metal and the soil is the electrolyte.

Fore more details on cathodic protection refer the following link:

http://en.wikipedia.org/wiki/Cathodic_protection

Cathodic protection is often used with coatings to protect piping. Regardless of the care used in coating and installing buried lines there will often be small pinholes in the coating. A cathodic protection system can protect against corrosion at these points and significantly extend the life of the piping.

Cathodic protection is applied to underground piping as a system. At every location where cathodically protected pipe leaves the soil (or water) it must be electrically isolated from the aboveground continuation of the line if the continuation is not part of the CP system. This is done with an insulating flange kit (or insulating union on small diameter pipe) that uses electrically insulating bolt sleeves, nut washers, and sealing gasket in a conventional flange makeup.

Cross country steel pipelines, and steel submarine piping are the principal users of cathodic protection.

This is all folks, as a brief synopsis on the subject of buried pipes and I look forward to comments on the subject from the members of Cheresources.

Regards,
Ankur.

Determining Diesel Transfer Pump Capacity For Diesel Engines And Generators

Sat, 08 Nov 2014 09:46:00 +0000

Dear All,

Often process engineers are asked to design a diesel storage and transfer system for diesel engines and diesel power generators. Where to start is often a question among process engineers. There has to be a design basis for designing such a system. Today's blog entry addresses the design basis for design of a diesel storage and transfer system along with a solved example for such a design. Readers of my blog are welcome to convert this information into an excel spreadsheet calculation and share.

Consider the following calculation example:

Inputs:
No. of Installed DGs: 3
No. of Working DGs: 2
Rating per DG: 1500 kW
Diesel Filling time: 1 hour (to Diesel day tank)
Transfer Pump Design Flow factor: 1.1
Diesel Day tank storage basis: 24 hours
Usable Day tank volume factor = 0.95
Main Diesel Storage Tank hold-up basis = 7 days (Note 1)

Calculations:
Diesel Flow per DG = 0.4 m³/h (from table above with vlookup function)
Total Diesel flow = 0.8 m³/h (0.4*2)
Day tank theoretical capacity per DG = 9.6 m³ (0.4*24)
Day tank working capacity per DG = 10.11 m³ (9.6 / 0.95)
Max. Diesel Consumed per day = 19.2 m³ (9.6*2) where 2 is the no. of working DGs
Max. Diesel Rate = 19.2 m³/h (19.2 / 1) where 1 is the Diesel filling time in hours
Diesel Pump Design Capacity = 21.12 m³/h (19.2*1.1)
Main Diesel Tank working Capacity = 134.4 m³ (19.2*7)

Notes:
1. For on-shore installations the total diesel storage is often considered for 7 days and for off-shore installation it is generally considered as 15 days.
2. The above example calculation considers a separate main storage tank and separate individual day tanks for individual DG sets. In many instances, the individual day tank for DGs may have a hold-up for 8 to 12 hours instead of 24 hours which would require 2 fillings per day for 12 hours hold-up and 3 fillings for 8 hours hold-up

Hope readers of my blog find this blog entry interesting and I look forward to their comments.

Regards,
Ankur

Attached Thumbnails

Selection And Sizing Of Marine Loading Arms For Petroleum (Black / White Oil) Products

Sat, 04 Oct 2014 09:22:00 +0000

Dear All,

Todya's blog entry provides a procedure for selection / sizing of marine loading arms

Vessel / Tanker Design Discharge Capacity (Figure 1)

Oil Tanker Manifold Diameters

Loading Arm size should not be larger than a maximum of two sizes of the the ships manifold diameter. A smaller size loading arm compared to the ship's manifold diameter is acceptable using a reducing spool piece.

Flow Rate Vs No. of Marine Loading Arms (Figure 2)
Based on limiting flow velocity of 11 m/s (36 ft/s)

Maximum Marine Arm Loading Flow Rates
Based on Limiting Flow Velocity of 11 m/s (36 ft/s)

Example Calculation

Inputs
Max Vessel Size = 100,000 DWT
Product Viscosity = Heavy
Pump Type = Centrifugal

Outputs
bbls/hr/DWT = 0.6 (From Figure1)
Correction factor = 0.88
Max. Discharge Rate = 52,800 bbls/hr
Max. Arm Size = 16 inch
No. of Arms = 2 (From Figure 2)

Note:
Part III, Section 2-Guidelines (page 40) of the document "Design & Construction Specification for Marine Loading Arms" which is issued as a standard for MLAs by the "Oil Comapnies and International Marine Forum" with the acronym "OCIMF" recommends a maximum velocity of 12 m/s through marine loading arms.

Looking forward to hear comments from the members of the "Cheresources" community.

Regards,
Ankur.

Attached Thumbnails

Flare Dispersion Analysis - Frequently Asked Questions (Faqs)

Mon, 08 Sep 2014 16:01:00 +0000

Q1. What is Flare Dispersion Analysis?
A1. Flare dispersion analysis is a study of the dispersion of toxic and / or flammable pollutants (primarily gaseous) from a flare, either a ground or elevated flare.

Q2. Is it necessary to perform a flare dispersion analysis?
A2. Yes, it has now become mandatory in most parts of the world when a new industrial unit which requires a flare to burn waste gases requires to be set up and an environmental clearance is required from the local government for establishing the unit.

Q3. Is flare dispersion study required for existing old industrial units?
A3. This depends on the local laws and regulations related to environmental safety for the concerned industrial unit. From a moral and ethical viewpoint, it becomes obligatory for any organization to conduct not only a flare dispersion study for a new industrial unit it owns but also for old industrial units which may be more prone to emitting harmful pollutants to the atmosphere. Changing demographics make a compelling reason for periodic flare dispersion studies over the lifetime of the industrial unit. Most advanced nations have now made it compulsory for such kind of studies whether it be a new industrial unit, license renewal for an existing old unit or due to the changing demographics around a under construction / existing industrial unit.

Q4. Which are the reputed statutory/regulatory and guiding agencies which provide guidelines and assistance in conducting flare dispersion studies?
A4. Refer the list below:
a. United States Environmental Protection Agency (EPA)
http://www.epa.gov/
b. European Environment Agency
http://www.eea.europa.eu/
c. Oil Industry Safety Directorate (OISD) - India
http://oisd.gov.in/
d. Ministry of the Environment – Government of Japan
http://www.env.go.jp/en/
e. Environment Agency – Gov. UK
https://www.gov.uk/government/organisations/environment-agency

Q5. What kind of dangerous pollutants are normally encountered from an industrial flare?
A5. Some of the most common pollutants are listed below:
a. SO₂ (Sulfur Dioxide)
b. NO₂ (Nitrogen oxide)
c. CO (Carbon monoxide)
d. CO₂ (Carbon Dioxide) (Greenhouse gas)
e. Unburnt Hydrocarbons
f. Suspended Particulate Matter (SPM)
g. H₂S (Hydrogen Sulfide)

Q6. What basic scenarios need to be considered when doing a flare dispersion analysis?
A6. Flares may operate under different conditions depending on the type of industrial unit where the flare is installed. However, the basic three scenarios that need to be considered for a flare dispersion study are as follows:
a. Normal flaring under normal plant / unit operating conditions
b. Emergency flaring under emergency or plant / unit upset conditions
c. Flameout when combustion is not taking place and cold venting from the flare occurs

Q7. Which is the most well known mathematical model used for flare dispersion studies?
A7. The most well known mathematical model widely used for air dispersion modeling is the “Gaussian Dispersion Model”. Refer the link below:
http://en.wikipedia.org/wiki/Atmospheric_dispersion_modeling
http://petrowiki.org/Flare_and_vent_disposal_systems

Q8. Which are the well known commercial software that can be utilized for flare dispersion studies?
A8. “PHAST” from DNV, “FlareSim” from Softbit Consultants, and “FRED” from Shell Global Solutions

Q9. Which other software can be used for flare modeling?
A.9 A freeware "SCREEN3" for point source (flare) modeling to estimate the ground level concentration can be found at the following link:
http://www.weblakes.com/products/screen/
Refer the links below for some other commonly used software:
http://www.weblakes.com/download/us_epa.html#emissions
http://www.ehssoftserve.com/air_airmod.htm

Look forward to comments and observations from the readers of my blog.

Regards,
Ankur.

Some Guidelines For Fixed Bed Ion Exchange Units And Forced Draft Degasifiers

Mon, 11 Aug 2014 17:36:30 +0000

Dear All,

Boiler feedwater is often obtained from raw water by ion-exchange process utilizing vertical cylindrical vessels which contain ion-exchange resins of various types. These resins remove dissolved salts from the water and replace them with Hydrogen (H⁺) and hydroxyl (OH^-) thus producing pure, mineral free boiler feed quality water.

Resins which replace postively charged ions such as Calcium (Ca⁺⁺), Magnesium (Mg⁺⁺), Sodium (Na⁺) etc. are called "cation" resins and they are regenerated with an acid (sulphuric or hydrochloric) to replace the hydrogen ions which were exchanged for the "cations". More often sulphuric acid is employed for regeneration since it is cheaper than hydrochloric acid and can be stored in carbon steel vessels at ambient temperatures. Hydrochloric acid is more efficient in regeneration than sulphuric acid but storage and handling is expensive and more safety concerns need to be addressed. HCl storage requires elastomer-lined steel vessels and may require fume exposure protection measures such as vent scrubbers due to its high volatility.

Resins which replace negatively charged ions such as Sulphate (SO₄^--), Chloride (Cl^-), Carbonate (CO₃^--) etc. are called anion resins and they are regenerated with caustic (NaOH) to replace the hydroxyl ions which were exchanged by the "anions".

There are weak and strong cation as well as weak and strong anion resins. In general, the weak resins are more efficient in terms of regeneration chemicals (acid or caustic) consumption. However, these resins are also limited in their capability to remove ions. For example, cation removal with Weak Acid Cation (WAC) resin is limited to the amount of Bicarbonate (HCO₃^--) alakalinity in the water; and anion removal is limited to sulfides, chlorides and nitrates in the acidic form when using Weak Base Anion (WBA) resin.

Some cations and anions will slip through the ion exchange system and this is called "leakage".The amount aof leakage is a function of several factors such as:
a. Influent water quality
b. Type of resin employed for the cation and anion exchangers
c. Regeneration chemical type and dosage

There are some design guidelines / parameters for fixed bed ion exchange (excluding packed beds) units and forced draft degasifiers which are as follows:

1. Recommended influent water quality to avoid plugging, fouling and deterioration of resins is provided in the attached table:

2. Flowrates for design of vessel diameter should be kept in the 2 - 15 gpm/ft² (18 gpm/ft² if uniform resin is used) (1.4 - 8.1 L/s/m²) range. The lower limit prevents channeling and the maximum rate allows higher service flows during those periods when other vessels are in regeneration. The maximum flowrate for mixed bed units in polishing service should not exceed 24 gpm/ft² (16.3 L/s/m²).

3. Maximum and minimum bed depths should be 84 in. (2100 mm) and 30 in. (750 mm), respectively.

4. The maximum practical vessel diameter to be considered should be 15 ft (4.6 m).

5. Freeboard (as measured from the top of the exhausted resin bed) should be as follows:
Cation Units – 0.75 x resin bed depth + 42 in. (1050 mm)
Anion Units – 1.0 x resign bed depth + 48 in. (1200 mm)
Mixed Bed Units – 1.0 x resin bed depth + 48 in. (1200 mm)

6. Units should be designed with a maximum of one regeneration per 24 hours.

7. Forced draft degasifiers (also known as Decarbonators) typically reduce free carbon dioxide (CO₂) to 5 ppm and oxygen (O₂) to the 6 - 8 ppm range. The following design guidelines should be considered:
a. Maximum allowable water flowrate should be 17.5 gpm/ft² (11.9 L/s/m²) of cross sectional tower area. Minimum tower height should be 16 ft (4800 mm).

b. The minimum allowable air flowrate should be 3 scfm per gpm (20 L/s per L/s) of design water flow.

c. Storage capacity should not be less than 10 minutes of design throughput rate. Additional storage capacity may be needed for regeneration water requirements.

Design of Ion-Exchange DM water plants is often vendor specific and manufacturers / vendors often provide performance guarantees for outlet DM water quality, regeneration chemicals consumption and efficiency in terms of cubic meter or gallon of Demineralized water produced per cubic meter or gallon of raw water feed to the DM water plant. However, the above mentioned guidelines can help process engineers in evaluating and making an informed judgement on the offers and specifications submitted by the vendors / manufacturers.

Hope readers of my blog find this blog entry informative and I look forward to comments from them.

Regards,
Ankur.

Attached Thumbnails

Dielectric Constant Data Required For Guided Wave Radar Type Level Transmitters

Tue, 22 Jul 2014 16:44:00 +0000

Dear All,

One of the most important data required for specifying guided wave radar level transmitters for storage tanks is the Dielectric constant of the stored liquid in the tank.

This data is provided by process engineers. A lot of this data is available directly form various open sources form the internet. Some links are provided below:

http://www.engineeringtoolbox.com/liquid-dielectric-constants-d_1263.html

http://krohne.com/en/services/dielectric-constants/

However, Dielectric constant can be calculated as a value using refractive index values of the particular liquid. Following is the correlation between refractive index and dielectric constant

Dielectric constant = (Refractive Index)²

or

Refractive Index = (Dielectric Constant)^o.5

It is important to note that the above correlation is good for non-polar liquids. As the liquid deviates from a non-polar behavior to a polar behavior, errors are observed in the values obtained.

Benzene and Toluene are examples of non-polar liquids whereas Glycerol is a polar liquid.

For petroleum distillates a quotient "I" is defined as the refractive index parameter and is related to the refractive index as follows:

I = (n² - 1) / (n² + 2) ---------- (1)

For petroleum fractions having an average molecular wieight below 300 the following correlation is given for the refractive index parameter:

I = 0.328*v^-0.003*SG^0.915---------- (2)

where:
v = Kinematic viscosity of the petroleum fraction, cSt
SG = Specific gravity of the petroleum fraction

Let us take an example for calculation of a diesel :

v = 5 cSt @storage temperature
SG = 0.85 @storage temperature

From eqn (2),

I = 0.28

From eqn (1),

(n² - 1) / (n² + 2) = 0.28

Evaluating "n" the refractive index from above:

n = 1.47

Dielectric constant is refractive index squared or "n²".

Hence,

Dielectric constant = 2.16 (~2.2)

Hope, process engineers who need to provide the dielectric constant for specifying "Guided Wave Radar Level Transmitters" will find this blog entry informative.

Looking forward to hear comments from the readers of my blog.

The reference for the aforementioned equations is provided below as a web link:

http://kosalmath.files.wordpress.com/2010/08/hydrocarbons.pdf

Regards,
Ankur

Rotary Screw Compressors - Discussion And Calculations

Thu, 03 Jul 2014 20:34:46 +0000

Dear All,
Rotary Screw compressors are fast replacing piston type reciprocating compressors all over the world. With advancement in technology of precision machining the helical screws required in screw compressors, they are becoming the first choice for selection where otherwise a piston reciprocating compressor would have been used in earlier days.

I had provided a chart for compressor selection in one of my posts on “Cheresources” at the following link:

http://www.cheresources.com/invision/topic/11569-compressor-selection/?p=43854#entry43854

While this chart is quite handy for compressor selection, latest advancements in technology of screw compressors allow far higher capacities and higher pressure ratios (discharge pressure / suction pressure) compared to what is shown in the chart.

Some general advantages of screw compressors in the category of positive displacement compressors are listed below point wise:

1. Less pulsations or surging of flow compared to piston type compressors since the gas compression process with a rotary screw is a continuous sweeping process.

2. Low mechanical vibrations similar to centrifugal machines compared to piston reciprocating compressors.

3. The inlet or suction volume flow and power consumption increase linearly with the compressor speed at constant discharge pressure.

4. Compared to a centrifugal compressor the inlet or suction volume is nearly constant for variation of pressure ratio or gas molecular weight with no surging limit.

5. The achievable pressure ratio per compressor stage is not limited by the gas molecular weight but limited only by the allowable discharge temperature and mechanical limits of the compressor due to high temperatures.

6. Very high pressure ratios up to 10 per compressor stage can be achieved by liquid injection for cooling.
The pioneers of Screw compressor technology were the Swedish company “Svenska Rotor Maskiner” abbreviated as SRM.

The history of the development of the screw compressor can be found at the following link:

http://www.opcon.se/web/History_4.aspx

SRM main web page can be found at the link provided below:

http://www.opcon.se/web/SRM_EN.aspx

With this blog entry I am also attaching an excel workbook which provides some design equations for calculating the shaft power of oil-free or dry screw compressors. The reference for the design equations is the book “Compressors – Selection and Sizing” by Royce N. Brown which I personally consider amongst the best books for compressors. The book provides equations in USC units which I have converted to SI units. The reference for the rotor diameters and L/d ratios is taken from the reference provided in the link below:

http://www.mandieselturbo.com/files/news/filesof12079/949e.pdf

Other diameters and L/d ratios used by various manufacturers may be utilized in the excel workbook.

I welcome comments on my blog entry from the knowledgeable members of the “Cheresources” community.

Quick note from the admin: You can download the MS Excel workbook that accompanies this blog entry in the File Library.

Regards,
Ankur.

Technical Aspects Of Equipment Procurement

Fri, 06 Jun 2014 20:05:00 +0000

Dear All,
One of the readers of my blog made a suggestion about writing something related to equipment procurement. Well, I haven’t exactly acted as a procurement engineer in my career but fortunately did have the opportunity to interact with a few procurement engineers and pick up some ideas related to the methodology of equipment procurement.

One important thing to note about equipment procurement is that this activity is peculiar to the geographical location of the plant or unit. What it means is that equipment procurement has its own individual characteristics depending on whether this is being done in North America, European Union, Middle East, Africa, Asia-Pacific or South Asia.

I will not dwell on the logistics of equipment procurement since I have very little knowledge on the subject. However, I can certainly share some ideas related to the technical aspects of equipment procurement. Some of these ideas are shared point wise below:

1. Either the client (the knowledgeable ones) or the EPC contractor has an “Approved Vendor” list. The procurement engineer is bound by this vendor list and has no need to look further beyond this list for floating an enquiry. In fact, having an approved vendor list makes the procurement activity simpler.

2. Some companies have a registration process for being an approved vendor for the company. This may cause the “Approved Vendor” list to be too big with numerous vendors. Floating an enquiry for a particular equipment to all the vendors given in the “Approved Vendor” list for that equipment would not be a very practical thing to do. An experienced procurement engineer would probably trim the list of these vendors and restrict his enquiry to a top few vendors based on his knowledge of the vendor’s capabilities. Personally, I would restrict floating an enquiry to a maximum of four (4) vendors.

3. An experienced procurement engineer would ensure that the enquiries for the “long-lead” items are fast-tracked. A general definition of “Long-lead” items is an item that takes a longer time to procure. Various companies have their own norms on deciding what is meant by long-lead items depending on the type of plant that is being installed. As a rule-of thumb, I would consider any item with a lead time of more than 8 calendar months to be long-lead in today’s perspective of fast-tracked plants.

4. A good procurement engineer would keep continuous track of the response provided by the vendors in terms of submission of offers as per the enquiry. He or she would also be familiar with the process of raising “technical queries” often abbreviated as “TQ” on the offers provided by the vendors. In fact a procurement engineer who has an engineering design background prior to his taking up the responsibilities as a procurement engineer would be able to prepare a standard TQ format for various type of standard equipment. Having a standard TQ format saves time in the overall procurement cycle.

5. TQ’s may be raised by individual engineering disciplines in their own TQ formats. The procurement engineer will require collating all this information in one document for sake of clarity.

6. Most of the times information flows via e-mails with attachments and occasional telephonic conversation and video conferencing. I am a firm believer in a face-to-face communication with vendors. The resolution of unanswered questions and issues takes place best when you are sitting across the table and discussing. This has probably to do with human psychology. Some of the good procurement engineers I have interacted follow this very principle that I am advocating. Obviously, a good and meaningful way to have a discussion sitting across the table is to provide an advance discussion agenda to the party who you want to discuss with.

7. Once the offers from the vendors are streamlined to match the requirements of the project based on TQ resolution, the next step of “Technical Bid Analysis” abbreviated as “TBA” is conducted by the engineering design disciplines with active support and follow-up from the procurement engineer. Most good engineering companies have standard TBA formats for standard process plant equipment. The purpose of the TBA is to compare the offers of the vendors in the race for bagging the equipment order strictly from a technical view-point. The standard TBA worksheet has the benchmarking parameters of the particular equipment in the first column of the worksheet and the subsequent columns are information provided by the vendors against those benchmarked parameters. As an example, if you have 3 vendors there will be 3 columns in the worksheet for filling the data provided by the vendors against the benchmarked parameters. A summary section will be provided in the worksheet which will rate the vendors overall technical capabilities and provide rankings and in certain cases even a rejection of one or more of the vendors.

8. The next step would be the commercial negotiations for the equipment order. As an example, if three vendors have provided offers and the TBA concludes that all three of them are equal in their technical capabilities then by conventional logic it boils down to selecting a vendor offering the lowest first cost of the equipment. Please note, that this may not always be true and other factors such as market reputation of the vendor, past association with the company and other extraneous factors may take precedence in selection of the vendor.

9. The last step from a technical view-point in the procurement cycle would be placing the order of the equipment to the selected vendor in the form of a “Purchase Order" (PO) specification or requisition along with the commercial terms and conditions of purchase. The PO specification would generally include some technical data and information specific to the selected vendor's evaluated and accepted offer.

The aforementioned steps, however, do not end the procurement cycle or the responsibility of the procurement engineer. The logistics of equipment procurement now come into the forefront including items such as stage-wise inspection, factory acceptance test, test certification, packing, shipping, warehouse storage, insurance and demurrage. I will not deal with these items due to my inherent lack of knowledge on these matters. Maybe some experienced procurement engineer can provide some insight about the logistics part of the procurement process.

I look forward to a lot of comments from the readers of my blog.

Regards,
Ankur.

Managing Technical Information Exchange With Process Equipment / Package Vendors

Thu, 15 May 2014 11:52:00 +0000

Dear All,

Today’s blog entry has shades of technical as well as communication skills. It is not restricted to process engineers and the guidelines provided in the subsequent paragraphs can be utilized by engineers from other disciplines such as Projects, Mechanical, Electrical, Instrumentation etc. in how to manage their dealings with manufacturers / vendors.

Before I begin, let us make a distinction between a manufacturer and vendor.

A manufacturer is the one who actually manufactures / fabricates and / or assembles the pieces of equipment for the particular end-use. A manufacturer need not necessarily assemble the pieces of equipment for end-use. This may be done by a third party who conforms to the assembly requirements of the equipment pieces as specified by the manufacturer. A common acronym which many experienced engineers are familiar with is “OEM” which expands to “Original Equipment Manufacturer”.

“Vendor” is a loosely used term and could mean a manufacturer, an assembler, a trader or the combination of all the three. Most experienced engineers in the process industry are more familiar with this term and I will be using this term rather than manufacturer.

Where does all this begin? It begins from the requirement of a particular equipment for a given chemical process plant and for a given chemical manufacturing process.

How does it appear as information in the first place? It gets represented in a “Process Flow Diagram” in the very first stage. The second stage involves qualifying the function of the equipment in terms such as exact type of equipment, capacity, power requirement (rotary equipment), material of construction, operating and design conditions (pressure / temperature), properties of the process material entering / leaving the equipment, compliance to statutory requirements (if any), HSE requirements and any special considerations.

The aforementioned information is generally captured as a “Process Datasheet” by the process engineer.

The third step would be to prepare a specification or material requisition which would provide further details for the requirement to be fulfilled. Some further information required to be part of a specification could be as follows:
- International / Company Standards that need to be followed in design and manufacture
- Vendor Documentation type and quantified in terms of number of paper / electronic documents
- Spares and consumables requirement in terms of commissioning and operation of the equipment
- Material testing and test certification requirement
- Inspection and testing of the completed equipment as per any established guidelines and standards.
- Packing and Shipping Instructions
- Site Conditions for equipment storage, erection and commissioning
- Erection / Commissioning / Operation / Maintenance instructions
- Lubrication Schedule

These are jus t a few to name. There could be many more depending on the type of equipment and specific project requirement.

Now let us discuss how to improve the communication with the vendor to expedite the procurement of the equipment as required or suited for the particular application and in a timely manner. Please note that the emphasis is on suitability of the equipment for a particular function and an adjective such as “best” has not been used to describe the equipment.

1. As far as possible avoid providing data as a wide range. Do your study to arrive at precise data and parameters for the equipment. There is practically no equipment which can perform well over a wide range. You are creating hurdles for the vendor to come up with a suitable equipment design by providing a very wide range of data. You should know the operating or rating case of your equipment.
2. Wherever possible please fill the information in the datasheet. Do a thorough analysis of what information you can provide and what you cannot. Ask vendor to provide only that information that is relevant for the equipment and is necessary for the selection of the equipment.

3. Asking for design information related to any proprietary design is a waste of time. It is better to prevail on the vendor to provide performance guarantees for a proprietary design. However, make sure by doing your homework that what the vendor is claiming to be a proprietary design, is really proprietary, and not something that is an open design.

4. Incomplete or incorrect information can be disastrous in terms of getting trapped in selecting the wrong equipment and allowing the vendor to escape from his contractual obligations for the performance guarantee of the equipment.

One of the most common examples of this would be the selection of a wastewater treatment plant. Incomplete or wrong chemical and biological analysis of the influent waste water would certainly lead to selection of the wrong treatment unit or units. This would be specifically true when the specifications of the discharged effluent from the treatment plant are very stringent and there are practically no margins on the discharged effluent quality as specified by regulatory authorities. It also gives the vendor the excuse of escaping from the performance guarantee clause of the contract by proving that the influent water quality is not correct or incomplete for his specific design. In a nutshell, wastewater analysis in terms of multiple samples over a spread of time would be the ideal way to ensure that the treatment plant selected gives the optimum performance in terms of quality of discharged effluent.

One major problem related to providing such data is how to manage this in case of a Grass-roots or Greenfield project. This is also manageable. Most reputed vendors maintain an extensive databank of projects they have executed in the past. It is very likely that they would have data related to a similar project executed in the past which could be used for your project. Here my emphasis is on reputed vendors with an extensive portfolio of executed projects.

5. Be prompt to raise queries to the vendor when you find an error in the vendor documents or something which does not fit into the scheme of your plant or unit. Remember the vendor is in the market to do business and earn a profit. You must understand your requirement precisely. Many times the vendor would try to sell you something extra which is not required for your application. The analogy goes like this:
“I only need a sedan. Why are you trying to sell me a limousine?”

I hope the above mentioned guidelines provide an insight to the new engineers on how to effectively communicate with vendors in order to ensure a proper and timely selection of any equipment or package.

I look forward to comments from the members of “Cheresources”.

Regards,
Ankur.

Gas Boot Sizing Upstream Of Fwko Tanks

Mon, 28 Apr 2014 05:44:00 +0000

Gas Boots are provided upstream of a Crude Oil Dehydration tank in order to degass the crude oil before it enters the dehydration tank. Entrained gas in crude oil disturbs the settling process in a dehydration tank and it's removal is required to ensure quick and effective settling for separating bulk water in crude oil dehydration tanks also referred commonly as "Free Water Knockout" (FWKO) tanks.

It is important to note that Gas Boots are very effective up to a Gas-Oil-Ratio (GOR) of 10 Nm³/m³. Above the aforementioned GOR, a conventional gas-liquid separator is recommended.

In a gas boot the downward liquid velocity should not exceeed 0.1 m/s.

Gas boots are generally designed as a piece of piping in the inlet piping to the FWKO just uptream of the inlet connection to the FWKO. Today's blog entry provides a quick sizing method in terms of the diameter of the gas boot pipe and the wet and unstabilized crude oil nominal flow rate based on a reputed company guideline.

Refer the attachment for the sizing of gas boots.

Any comments are welcome.

Regards,
Ankur.

Attached Thumbnails

An Introduction To Corrosion Inhibitors

Tue, 04 Mar 2014 20:45:12 +0000

Dear All,

There have been quite a few threads devoted to injection of corrosion inhibiting chemicals on "Cheresources". Today's blog entry makes an attempt to give some introductory knowledge related to "Corrosion Inhibitors".

The following paragraphs provide the details. The reference for the content of this blog article is provided at the end of the blog article:

The use of chemical inhibitors to decrease the rate of corrosion processes is quite varied. In the oil extraction and processing industries, inhibitors have always been considered to be the first line of defense against corrosion. A great number of scientific studies have been devoted to the subject of corrosion inhibitors. However, most of what is known has grown from trial and error experiments, both in the laboratories and in the field. Rules, equations, and theories to guide inhibitor development or use are very limited. By definition, a corrosion inhibitor is a chemical substance that, when added in small concentration to an environment, effectively decreases the corrosion rate. The efficiency of an inhibitor can be expressed by a measure of this improvement:

Inhibitor Efficiency (%) = 100*(CR_uninhibited -CR_inhibited) / CR_uninhibited
where:
CR_uninhibited = corrosion rate of the uninhibited system
CR_inhibited = corrosion rate of the inhibited system

In general, the efficiency of an inhibitor increases with an increase in inhibitor concentration (e.g., a typically good inhibitor would give 95% inhibition at a concentration of 0.008% and 90% at a concentration of 0.004%). A synergism, or cooperation, is often present between different inhibitors and the environment being controlled, and mixtures are the usual choice in commercial formulations. The scientific and technical corrosion literature has descriptions and lists of numerous chemical compounds that exhibit inhibitive properties. Of these, only very few are actually used in practice. This is partly because the desirable properties of an inhibitor usually extend beyond those simply related to metal protection. Considerations of cost, toxicity, availability, and environmental friendliness are of considerable importance.

Attached Table presents some inhibitors that have been used with success in typical corrosive environments to protect the metallic elements of industrial systems. Commercial inhibitors are available under various trade names and labels that usually provide little or no information about their chemical composition. It is sometimes very difficult to distinguish between products from different sources because they may contain the same basic anti-corrosion agent. Commercial formulations generally consist of one or more inhibitor compounds with other additives such as surfactants, film enhancers, de-emulsifiers, oxygen scavengers, and so forth. The inhibitor solvent package used can be critical in respect to the solubility/dispersibility characteristics and hence the application and performance of the products.

Classification of Inhibitors
Inhibitors are chemicals that react with a metallic surface, or the environment this surface is exposed to, giving the surface a certain level of protection. Inhibitors often work by adsorbing themselves on the metallic surface, protecting the metallic surface by forming a film. Inhibitors are normally distributed from a solution or dispersion. Some are included in a protective coating formulation. Inhibitors slow corrosion processes by:

- Increasing the anodic or cathodic polarization behavior (Tafel slopes)
- Reducing the movement or diffusion of ions to the metallic surface
- Increasing the electrical resistance of the metallic surface

Inhibitors have been classified differently by various authors. Some authors prefer to group inhibitors by their chemical functionality, as follows:

Inorganic inhibitors: Usually crystalline salts such as sodium chromate, phosphate, or molybdate. Only the negative anions of these compounds are involved in reducing metal corrosion. When zinc is used instead of sodium, the zinc cation can add some beneficial effect. These zinc-added compounds are called mixed-charge inhibitors.

Organic anionic: Sodium sulfonates, phosphonates, or mercaptobenzotriazole (MBT) are used commonly in cooling waters and antifreeze solutions.

Organic cationic: In their concentrated forms, these are either liquids or waxlike solids. Their active portions are generally large aliphatic or aromatic compounds with positively charged amine groups.

However, by far the most popular organization scheme consists of regrouping corrosion inhibitors in a functionality scheme as follows:

Passivating (anodic)
Passivating inhibitors cause a large anodic shift of the corrosion potential, forcing the metallic surface into the passivation range. There are two types of passivating inhibitors: oxidizing anions, such as chromate, nitrite, and nitrate, that can passivate steel in the absence of oxygen and the non-oxidizing ions, such as phosphate, tungstate, and molybdate, that require the presence of oxygen to passivate steel.

These inhibitors are the most effective and consequently the most widely used. Chromate-based inhibitors are the least-expensive inhibitors and were used until recently in a variety of application (e.g., recirculation-cooling systems of internal combustion engines, rectifiers, refrigeration units, and cooling towers). Sodium chromate, typically in concentrations of 0.04 to 0.1%, was used for these applications. At higher temperatures or in fresh water with chloride concentrations above 10 ppm higher concentrations are required. If necessary, sodium hydroxide is added to adjust the pH to a range of 7.5 to 9.5. If the concentration of chromate falls below a concentration of 0.016%, corrosion will be accelerated. Therefore, it is essential that periodic colorimetric analysis be conducted to prevent this from occurring. In general, passivation inhibitors can actually cause pitting and accelerate corrosion when concentrations fall below minimum limits. For this reason it is essential that monitoring of the inhibitor concentration be performed.

Cathodic
Cathodic inhibitors either slow the cathodic reaction itself or selectively precipitate on cathodic areas to increase the surface impedance and limit the diffusion of reducible species to these areas. Cathodic inhibitors can provide inhibition by three different mechanisms: (1) as cathodic poisons, (2) as cathodic precipitates, and (3) as oxygen scavengers. Some cathodic inhibitors, such as compounds of arsenic and antimony, work by making the recombination and discharge of hydrogen more difficult. Other cathodic inhibitors, ions such as calcium, zinc, or magnesium, may be precipitated as oxides to form a protective layer on the metal. Oxygen scavengers help to inhibit corrosion by preventing the cathodic depolarization caused by oxygen. The most commonly used oxygen scavenger at ambient temperature is probably sodium sulfite (Na₂SO₃).

Organic
Both anodic and cathodic effects are sometimes observed in the presence of organic inhibitors, but as a general rule, organic inhibitors affect the entire surface of a corroding metal when present in sufficient concentration. Organic inhibitors, usually designated as film-forming, protect the metal by forming a hydrophobic film on the metal surface. Their effectiveness depends on the chemical composition, their molecular structure, and their affinities for the metal surface. Because film formation is an adsorption process, the temperature and pressure in the system are important factors. Organic inhibitors will be adsorbed according to the ionic charge of the inhibitor and the charge on the surface. Cationic inhibitors, such as amines, or anionic inhibitors, such as sulfonates, will be adsorbed preferentially depending on whether the metal is charged negatively or positively. The strength of the adsorption bond is the dominant factor for soluble organic inhibitors.

These materials build up a protective film of adsorbed molecules on the metal surface, which provides a barrier to the dissolution of the metal in the electrolyte. Because the metal surface covered is proportional to the inhibitor concentrates, the concentration of the inhibitor in the medium is critical. For any specific inhibitor in any given medium there is an optimal concentration. For example, a concentration of 0.05% sodium benzoate or 0.2% sodium cinnamate is effective in water with a pH of 7.5 and containing either 17 ppm sodium chloride or 0.5% by weight of ethyl octanol. The corrosion due to ethylene glycol cooling water systems can be controlled by the use of ethanolamine as an inhibitor.

Precipitation inhibitors
Precipitation-inducing inhibitors are film-forming compounds that have a general action over the metal surface, blocking both anodic and cathodic sites indirectly. Precipitation inhibitors are compounds that cause the formation of precipitates on the surface of the metal, thereby providing a protective film. Hard water that is high in calcium and magnesium is less corrosive than soft water because of the tendency of the salts in the hard water to precipitate on the surface of the metal and form a protective film.

The most common inhibitors of this category are the silicates and the phosphates. Sodium silicate, for example, is used in many domestic water softeners to prevent the occurrence of rust water. In aerated hot water systems, sodium silicate protects steel, copper, and brass. However, protection is not always reliable and depends heavily on pH and a saturation index that depends on water composition and temperature. Phosphates also require oxygen for effective inhibition. Silicates and phosphates do not afford the degree of protection provided by chromates and nitrites; however, they are very useful in situations where nontoxic additives are required.

Volatile corrosion inhibitors
Volatile corrosion inhibitors (VCIs), also called vapor phase inhibitors (VPIs), are compounds transported in a closed environment to the site of corrosion by volatilization from a source. In boilers, volatile basic compounds, such as morpholine or hydrazine, are transported with steam to prevent corrosion in condenser tubes by neutralizing acidic carbon dioxide or by shifting surface pH toward less acidic and corrosive values. In closed vapor spaces, such as shipping containers, volatile solids such as salts of dicyclohexylamine, cyclohexylamine, and hexamethylene-amine are used. On contact with the metal surface, the vapor of these salts condenses and is hydrolyzed by any moisture to liberate protective ions. It is desirable, for an efficient VCI, to provide inhibition rapidly and to last for long periods. Both qualities depend on the volatility of these compounds, fast action wanting high volatility, whereas enduring protection requires low volatility.

I would enjoy having a discussion and some insight on "corrosion inhibitors" from knowledgeable members of the forum.

Reference: Chapter 10, Corrosion Inhibitors" "Handbook of Corrosion Engineering" by Pierre R. Roberge

Regards,
Ankur

Attached Thumbnails

Uklpg (British) Code Of Practice For Safety In Bulk Lpg Storage

Tue, 04 Feb 2014 07:39:00 +0000

Dear All,

LPG storage and the related safety considerations have been discussed many times on various threads on 'Cheresources'. Any reader / member of 'Cheresources' can find these threads using the search facility on 'Cheresources'.

Today's blog entry presents the British perspective on safety related to bulk LPG storage. The reference for this blog entry has been provided at the end of this blog entry. Let us understand what the British view on safe LPG bulk storage means:

Safe storage and handling of LPG is of paramount importance, whether it is in bulk or in cylinders. This is basically achieved by ensuring the mechanical integrity of the storage vessels or cylinders and by strict observance of the recommended separation distances between storage and buildings or boundaries. This passive protection has to be supplemented by rigorous observance of operational procedures. The Liquefied Petroleum Gas Association (LPGA) has issued a Code of Practice on bulk LPG storage at fixed installations in collaboration with and approval of the Health and Safety Executive. Part 1 deals with design, installation and operation of vessels above ground.

Note: UKLPG was formed by the merger of the LP Gas Association (LPGA) and the Association for Liquid Gas Equipment and Distributors (ALGED) in January 2008. Its roots are firmly established, with LPGA and ALGED established in 1947 and 1975 respectively.

Perhaps the most serious hazard to LPG storage is that of accidental fire. The safety distances are intended to separate the storage from possible adjacent fires so that the risk of a fire affecting the storage is very low. However, this residual risk has to be catered for. The mechanical integrity needs to be assured under severe fire attack. For this reason, the vessels are provided with relief valves designed to cope with fire engulfment. The heat from such a fire may raise the stored pressure until the relief valve opens. The discharge capacity of the relief valve when fully open is required to meet or exceed the following:

Q = 10.6552*A^0.82

where:
Q = Relief Valve capacity in m³/min of air @15°C, 1.01325 bar (abs)
A = total surface area of the vessel, m²

The heat transfer to the liquid from an engulfment fire has been estimated at around 100 kW/m², and the above formula equates this to the vapor produced from this input as latent heat. The exponential is an area exposure factor which recognizes that large vessels are less likely to be completely exposed to flames.

The safety relief valve will protect the liquid-wetted areas of the storage vessel. The metal temperature will not significantly exceed the liquid temperature which will be absorbing the latent heat of vaporization. However, above the liquid line no such cooling will take place. The metal temperature at the top of the vessel could therefore exceed safe limits.

The usual protection for large installations is to provide a water-spray system. For small bulk storage, fire hoses or monitors are often adequate. However, for installations over 50 tons of storage (and all major cylinder-filling plants) it is accepted that a fixed water-spray system needs to be provided which is automatically initiated by a system capable of detecting a fire threatening the vessels and/or the adjacent tanker loading or unloading area. The deluge rate to provide protection against fire engulfment is 9.8 liter/min/m² of vessel surface area, and this should be capable of being sustained for at least 60 minutes. The spray pattern adopted for a fixed installation normally includes four longitudinal spray bars, two at the upper quadrants and two at the lower quadrants of horizontal cylindrical vessels, with nozzles spaced to give uniform coverage.

An alternative means of avoiding the hazard from fire is to bury the vessels or to employ the increasingly popular method of mounding. In either case acknowledgment of the reduced hazard is indicated by the reduced separation distances (see attached table). Since either burial or mounding preclude the possibility to monitor continuously the external condition of the vessels, very high-quality corrosion protection needs to be applied, normally supplemented by cathodic protection, depending on soil conditions. Guidance approved by the H.S.E. is published by the LPGA in their Code of Practice 1 part 4.

The use of burial or mounding is sometimes employed to overcome visual environmental objections since the mounding, for instance, can be grassed over. Indeed, this is the method often adopted to prevent erosion of the mounding material.

Hope all of you have enjoyed this small article related to safety on Bulk LPG storage as followed in the UK.
I do expect some comments and will be glad to share my views on the comments.

Regards,
Ankur

Reference: Chapter 14, Plant Engineer's Reference Book , 2^nd Edition by Dennis A. Snow

Attached Thumbnails

Design Equations For Venturi Scrubbers

Mon, 27 Jan 2014 04:52:00 +0000

Dear All,

Many of you involved in engineering and design of wet scrubbing systems are already familiar with “Venturi Scrubbers”.

Venturi Scrubbers are primarily designed to control fine particulate matter (PM) from gaseous streams. Before releasing waste gases that contain PM to the atmosphere, these gases are treated using venturi scrubbers to remove PM. Generally the effective range of PM removal is from 10 microns to 2.5 microns. They are also capable of some incidental control of “Volatile Organic Compounds”, which, however is not their primary function. They can also be used for capturing high solubility gases which have good solubility with the sprayed liquid.

Venturi scrubbers have high efficiencies when capturing PM in the range of 0.5-5 micron.

Venturi scrubbers have following typical industrial applications:
a. Boiler waste gases utilizing coal, oil, biomass and liquid waste
b. Metal Processing – Iron & Steel, Aluminum
c. Wood, Pulp & Paper Industry
d. Chemical Industries
e. Municipal Solid Waste Incinerators

For a detailed description on venturi scrubbers refer the Wikipedia link below:
http://en.wikipedia.org/wiki/Venturi_scrubber

Today’s blog entry relates to some design equations for evaluating liquid droplet diameter, collection efficiency, throat velocity, throat diameter, throat length and pressure drop for venturi scrubbers:

Liquid Mean Droplet Size or Sauter Mean Diameter

Nukiyama & Tanasawa Correlation

d_l = (0.000585/v_r)*sqrt(σ/ρ_l) + 0.0597*(µ_l /sqrt(σ/ρ_l))^0.45*(Q_l/Q_g)^1.5
where:
d_l = mean droplet diameter, m
v_r = relative velocity of gas to liquid, m/s = v_g –v_l ≈v_g
Note: In most cases, the gas velocity is much higher than the liquid velocity and v_r may be considered equal to v_g
σ = liquid surface tension, N/m
ρ_l = liquid density, kg/m³
µ_l = liquid viscosity, Pa.s
Q_l = volumetric flow rate of liquid, m³/s
Q_g = volumetric flow rate of liquid, m³/s

Boll et. al Correlation

d_l = (0.042 +0.00565*(1000*Q_l / Q_g)) / v_r^1.602

Collection Efficiency

η = 1 – e^(-k*R*sqrt(ψ))----(1)
where:
η = collection efficiency of the venture scrubber, fraction
k = correlation coefficient whose value depends on system geometry and operating conditions (typically 0.1-0.2)
R = liquid-to-gas ratio, m³/1000 m³
ψ = inertial impaction parameter, dimensionless
Note: R values between 0.936 m³/1000 m³ and 1.337 m³/1000 m³ provide optimum collection efficiency
ψ = C*d_p^2*ρ_p*v_t / (9*µ_g*d_l)-----(2)

where:
C = Cunningham Slip correction factor, dimensionless

C = 1 + (0.000621*Tg / (d_p*10^6))-----(3)
T_g = inlet gas absolute temperature, K
d_p = particle diameter, m
ρ_p = particle density, kg/m³
v_t = throat velocity, m/s
µ_g = gas viscosity, Pa.s
d_l = liquid mean droplet diameter, m

Normally collection efficiency is an input, so re-writing equation (1) in terms of ψ:
ψ = (ln(1-η) /(k*R))^2-----(4)

Since we want to know the throat velocity, re-writing equation (2) in terms of v_t:
v_t = ψ*9*µ_g*d_l / (C*d_p^2*ρ_p)-----(5)

Throat Length

l_t = 369.561*R^0.293 / v_t^1.127
where:
l_t = throat length, m
R = liquid-to-gas ratio in L/m³ (to convert m³/1000 m³ to L/m³ multiply m³/1000 m³ with 0.001)
v_t = throat velocity, m/s

Throat Area

A_t = Q_g / v_t
where:
A_t = throat area, m²
Q_g = process gas flow rate, m³/s
v_t = throat velocity, m/s

Pressure Drop in Venturi Scrubbers (Hesketh Equation)

ΔP = 0.532*v_t^2*ρ_g*A_t^0.133*(0.56 + 16.6*(Q_l/Q_g) + 40.7*(Q_l/Q_g)^2)
where:
ΔP = Pressure drop, Pa
v_t = throat velocity, m/s
ρ_g = gas density downstream of throat, kg/m³
A_t = throat area, m²
Q_l = volumetric flow rate of liquid, m³/s
Q_g = volumetric flow rate of gas, m³/s

Hope the readers of this blog entry have now some idea about the design equations related to venturi scrubbers. Please note that the liquid-to-gas ratios are basically ratios and any set of volumetric flow rate units may be used as long as they are consistent for both liquid and gas.

The entire blog entry has been a compilation from various resources related to Venturi scrubbers. However the following resources can be referenced from the links provided below:

http://web2.clarkson.edu/class/ch434/files/Omara%20and%20Dev_120408.pdf

http://tean.teikoz.gr/en/research/LAFEC/publications/publications/12_3.pdf

http://books.google.com.om/books?id=P1diAgAAQBAJ&pg=PA344&lpg=PA344&dq=pressure+drop+in+venturi+scrubber+Pa&source=bl&ots=29Quq4g-Gz&sig=zrIiXTURyvYHbugYCUrlUIA3hEg&hl=en&sa=X&ei=sN3kUpPXMI210QXBoYDAAQ&redir_esc=y#v=onepage&q=pressure%20drop%20in%20venturi%20scrubber%20Pa&f=false

I will try my best to provide answers to any questions raised. All these equations are programmable in an excel spreadsheet.

Regards,
Ankur.

Hyperfocal Distance Calculator

Tue, 07 Jan 2014 12:28:00 +0000

Dear All,

This blog entry is not related to chemical engineering but it certainly is related to the science of optics and should provide interesting reading for those chemical engineers who happen to be photography enthusiasts and want to know some intricacies related to the science of photography

I have always loved photography right from the days when I bought my first Konica Minolta film camera in 1987. However, for reasons unknown I was not able to pursue this hobby of mine in a serious manner. Recently, I rediscovered my lost passion for photography and have now started seriously pursuing it. To regain my lost touch I started reading quite a lot of material on photography and found some really interesting and enlightening reading material as a photography enthusiast.

It is quite possible that a lot many of the chemical engineers who are members and readers of "Cheresources" may be serious enthusiasts of photography and spend some of their spare time pursuing photography.

Let me also explain the difference between a casual photographer and an enthusiast photographer. A casual photographer uses the "Automatic" mode on his digital camera to click pictures and leaves it at that. An enthusiast photographer uses or tries to use the "semi-automatic" and "manual" modes on his camera to click pictures where his understanding and judgement of how he is clicking a particular scene or person comes into play. In a nutshell, an enthusiast photographer is concerned about the quality and the ability to draw serious attention to the picture he has clicked.

I love landscape photography and try to do it whenever I can go on a holiday to a tourist destination.

Optics in itself is a vast subject of physics. Modern digital cameras have evolved so much that today you can take pictures of stunning beauty and clarity with just a click of the shutter button. However, till date no amount of advancement in the science of optics has been able to mimic the functions of the human eye and its coordination with the brain.

I don't have the luxury to indulge myself in a detailed discourse on photography and optics. Suffice to say that those who have some knowledge of photography and specifically landscape photography will be able to grasp what I am trying to explain.

When you do any kind of photography you have to focus on an object in your camera viewfinder or LCD screen. The area in front of the object of focus is called as 'foreground' and the area behind the object of focus is called 'background'. Most of the times when an object is focused on the camera and a picture is taken either the foreground or background do not come out sharp in the picture. The objective of producing a good landscape photograph is to have everything in the frame to be sharp and not blurred, the exception being portrait photography.

Then hyperfocal distance determines how sharp your landscape photograph will be, if you focus your camera at a pre-designated distance from the camera lens.

In technical terms hyperfocal distance is defined as follows:

The hyperfocal distance is the closest distance at which a camera lens can be focused while keeping objects at infinity acceptably sharp. When the camera lens is focused at this distance, all objects at distances from half of the hyperfocal distance out to infinity will be acceptably sharp.
http://en.wikipedia.org/wiki/Hyperfocal_distance

How to calculate the hyperfocal distance? If you know or set the focal length of the lens of your camera and the aperture setting in say aperture priority or manual mode when you want to click the picture you can calculate the distance at which you should focus to get a very sharp landscape picture with both the foreground and the background being sharp.

The attachment with this blog entry calculates the hyperfocal distance for various digital camera brands based on a camera specific attribute called 'Circle of Confusion'.

Hoping to get comments from chemical engineers who are also photography enthusiasts.

Regards,
Ankur

Attached Files

Hyperfocal Distance Calculator.xlsx (9.47MB)
downloads: 456

Time Dependent Gas Release Through A Hole From A Pressurized Container

Mon, 23 Dec 2013 18:47:05 +0000

Dear All,

One of the very important books related to process safety is popularly known as the "Yellow Book" published from the Netherlands and having the full title as:

"Methods for the calculation of physical effects - due to releases of hazardous materials (liquids and gases)" - 'Yellow Book' - CPR14E

The best part of this book is that it is absolutely free to download from the internet. The units used in the book are purely SI units which might prove somewhat of a dampener fro those quite used to engineering units and English units. In its treatment of the physical phenomena of gaseous and liquid releases it is an absolute delight to read. The coverage is extensive and wherever possible numerical methods to enunciate the physics involved in releases is provided.

It also encompasses other safety related topics such as vapor cloud dispersion, vapor cloud explosion, heat flux from fires etc.

Below a link is provided for a free download of the book:

http://content.publicatiereeksgevaarlijkestoffen.nl/documents/PGS2/PGS2-1997-v0.1-physical-effects.pdf

The title of my blog entry does not suggest anything about the book but I wanted to make it clear that the topic I have chosen i.e. 'Time Dependent Gas Release through a hole from a pressurized container' is basically picked up from the 'Yellow Book'.

I have prepared a spreadsheet based on the theory and equations provided in the 'Yellow Book' for the subject with an example. A few relevant pages of the 'Yellow Book' with certain sections highlighted have also been provided in the spreadsheet.

Those process engineers who have access to HYSYS or any other similar process simulation software where vessel depressurization dynamic utility is available, may find that the results are not closely matching with the spreadsheet that has been presented along with this blog entry. The reason is that a few simplified assumptions have been done which the book also follows to arrive at the results. Some of these are as follows:
1. The process of the vessel depressurization is considered to be an adiabatic process only.
2. The specific heat or heat capacity at constant volume (C_v) is considered to be constant along the depressurization path although the dynamic change in pressure and temperature during the depressurization will have some effect on the specific heat value. In other words, there will be minor changes in the specific heat value at different dynamic values of pressure and temperature.
3. The specific heat ratio (C_p/C_v) is also considered to be constant along the depressurization path based on the same premise presented above in point number 2.

Again the idea of presenting a spreadsheet is to ensure that those process engineers who do not have access to expensive simulation software are not left out.

Hope all of you enjoy this blog post and the accompanying spreadsheet. I look forward to your comments and definitely look forward to comparisons of the example given in the spreadsheet with the results obtained form a "Dynamic Depressuring Utility" from a simulation software such as HYSYS or similar.

Regards,
Ankur.

Quick note from the admin:
Download the MS Excel sheet here:
http://www.cheresources.com/invision/files/file/308-time-dependent-gas-leak-outflow-through-a-hole/

Frequently Asked Questions (Faq) About Liquefied Natural Gas (Lng)

Tue, 03 Dec 2013 13:37:00 +0000

Dear All,

LNG is a subject of great interest these days considering that it has become a primary source of energy for many countries deficient in fossil fuel based energy resources. It is also of interest because it is among the cleanest of all fossil fuel based energy resources when compared with other fossil fuel energy sources such as coal and crude petroleum and its derivatives.

Today's blog entry tries to put in perspective what LNG is all about. The focus is on introducing what LNG is to new engineers starting their career. The blog entry has been made in the form of 'frequently asked questions' related to LNG. Details are excluded and certain web references are provided as hyper links which can guide the readers where to look for more details. Here goes the FAQ:

Q1. What is LNG?
A1. Liquefied Natural Gas (LNG) is natural gas (NG) that is cooled to the point that it condenses to a liquid.

Q2. What are the conditions for natural gas to be converted to LNG?
A2. Natural gas condenses to a liquid i.e. becomes LNG at a temperature of approximately -161°C (-256°F) at atmospheric pressure (101.325 kPaa).

Q3. Why is it required to convert natural gas to LNG?
A3. Liquefaction of natural gas to LNG reduces the volume of natural gas by approximately 600 times thus making it more economical to store natural gas and to transport gas over long distances for which pipelines are not economically feasible or there are other constraints.

Q4. How is LNG stored?
A4. LNG is stored in double-walled storage tanks at atmospheric pressure. The storage tank is really a tank within a tank. The annular space between the two tank walls is filled with insulation. Refer the link below for more information on LNG storage tanks:
http://en.wikipedia.org/wiki/LNG_storage_tank

Q5. What materials are used to construct LNG storage tanks?
A5. The inner tank in contact with the LNG is made of materials suitable for cryogenic service and structural loading of LNG. These materials include 9 percent Nickel Steel, aluminum and pre-stressed concrete. The outer tank is generally made of carbon steel or pre-stressed concrete.

Q6. How is LNG transported by ships?
A6. LNG tankers are double-hulled ships specially designed to prevent leakage or rupture in an accident. The LNG is stored in a special containment system within the inner hull where it is kept at atmospheric pressure and cryogenic temperature (-161°C). For more details refer the following link:
http://en.wikipedia.org/wiki/LNG_carrier

Q7. What is re-gasification of LNG?
A7. LNG has to be converted back to natural gas by warming in a controlled environment at the receiving and re-gasification terminal. The LNG can be warmed by passage through pipes heated by direct-fired heaters, or pipes warmed by seawater, or through pipes that are in heated water.

Q8. How is LNG quantified for trading (selling / buying) in the world market?
A8. LNG is generally quantified for trading on a mass basis in terms of millions of tons

Q9. Is there an approximate conversion for LNG mass to NG volume at standard conditions of 1 atmosphere pressure and 60°F?
A9. The approximate conversion for LNG mass to NG volume at standard conditions of 1 atmosphere pressure (14.7 psia /101.3 kPaa) and 60°F is as follows:

1 Metric ton of LNG or 1000 kg of LNG = 48,700 Scf of NG = 1379 Sm³ of NG

Note: Exact conversion depends on natural gas molecular weight

Q10. Is LNG flammable?
A10. LNG when suddenly released from its containment to atmosphere at ambient temperature conditions flashes and forms a vapor-air mixture. This vapor-air mixture forms a visible vapor-air cloud. As the vapor cloud gets warmer in contact with the ambient air it gets lighter than air and rises. A flammable vapor-air mixture can only form if the natural gas vapor concentration in the vapor-air mixture is between 5% and 15% by volume of the vapor-air mixture. A vapor-air mixture with less than 5% by volume of NG cannot cause a flame or fire since the concentration of the vapor (NG) in the vapor-air mixture is not enough to start or sustain a flame or fire or in other words the vapor-air mixture is too lean. The value 5% by volume is thus called the Lower Flammability Limit (LFL) for natural gas. Similarly a vapor-air mixture with more than 15% by volume of NG cannot cause a flame or fire since the concentration of the air (oxygen) in the vapor-air mixture is not enough to start or sustain a flame or fire or in other words the vapor-air mixture is too rich. The value 15% by volume is thus called the Upper Flammability Limit (UFL) for natural gas.

Q11. Is LNG explosive?
A11. LNG in its liquid form is not explosive. When LNG forms a vapor-air mixture and is in a confined space with no means to disperse it can cause a vapor cloud explosion if exposed to a source of ignition. Again for the vapor-air mixture to explode, the vapor concentration of natural gas in the vapor-air mixture has to be between 5% and 15% by volume. Thus the term LFL / UFL is interchangeably used with the term Lower Explosive Limit (LEL) and Upper Explosive Limit (UEL).

Q12. What are the main uses of LNG?
A12. LNG after re-gasification to NG is primarily used for power generation, home heating and as cooking gas.

For more details related to LNG in general refer to the links below:
http://www.beg.utexas.edu/energyecon/lng/
http://en.wikipedia.org/wiki/Liquefied_natural_gas

I have no claim whatsoever as an expert in the LNG field. This is a general information that has been provided in this blog. I will take questions from the readers and try to answer them to the best of my ability.

Happy reading.

Regards,
Ankur.

Pressure Drop Of Pseudo-Plastic Fluids In Pipes

Fri, 15 Nov 2013 02:48:00 +0000

Dear All,

Pseudo-plastic fluids are a type of non-newtonian fluids and account for a majority of the non-newtonian fluids commonly found in the chemical process industry and specifically in the food processing industry.

Today's blog entry provides a methodology to calculate the pressure drop of a pseudo-plastic fluid in a pipe. An example problem is solved at the end of the outlined method. Readers can use the example problem to develop a spreadsheet based calculation

For Pseudo-plastic non-newtonian fluids, the power law model is used to define the shear stress versus velocity gradient relationship. The power law equation is as follows:

τ = K*(dV/dy)ⁿ

^where:
τ = shear stress at distance y from pipe wall
K = flow consistency index
dV/dy = velocity gradient or shear rate
n = flow behavior index

For pseudo-plastic fluids n < 1, and for dilatant fluids n > 1. Since the apparent viscosity μ is the slope of the shear stress τ versus velocity gradient plot, we can calculate the apparent viscosity of a non-newtonian fluid that follows the power law from the following equation

µ =K*(dV/dy)^n-1

^or
K = 0.001*µ*(dV/dy)^1-n

where:
µ = Apparent Viscosity, cP
K = Flow consistency index, Pa.sⁿ
dV/dy = velocity gradient, s^-1
n = flow behavior index, dimensionless

Power-Law Reynolds Number:

Re_PL = 2^3-n*(n / (3n+1))ⁿ*((V^2-n*Dⁿ*ρ) / K)

where:
Re_PL = Power-Law Reynolds Number, dimensionless
V = Velocity of flow in pipe, m/s
D = Inside diameter of pipe, m
ρ = density of the fluid, kg/m³

Determining fanning friction factor 'f' for laminar flow of power-law fluids:

f = 16 / Re_PL

where:
f = fanning friction factor, dimensionless

Note:
The transition from Laminar to Turbulent flow for pseudo-plastic fluids is considered at a Reynolds number of 2100.

Determining fanning friction factor 'f' for turbulent flow of power-law fluids:

Refer attached chart

Friction Loss (in head) using Fanning friction factor:

h_f = 2*f*V²*L / (g*d)
where:
h_f = friction loss (in head) in the pipe, m (meters of liquid column)
L = Length of the pipe, m
g = Acceleration due to gravity, 9.81 m/s²

Example Problem:

A coal slurry is to be transported by horizontal pipeline. It has been determined that the slurry may be described by the power law model with a flow behavior index of 0.4, an apparent viscosity of 50 cP at a shear rate of 100 /s, and a density of 1442 kg/m³. What would be the pressure drop and the horsepower (HP) required to pump the slurry at a rate of 0.06 m³/s through an 8 in. Schedule 40 pipe that is 80 km long ?

Inputs:

n = 0.4
µ = 50 cP
dV/dy = 100 s^-1
ρ = 1442 kg/m³
D = 0.202 m (based on nominal pipe size in inches and Pipe schedule)
L = 80 km
Q = 0.06 m³/s

Results:
L = 80,000 m
K = 0.792 Pa.sⁿ
V = 1.87 m/s
Re_PL = 8025
f = 0.0048 (from chart)
h_f = 1358 m
ΔP = 192 bar
HP = 1153038 W = 1153 kW

Hope all of you have found this blog entry very informative. I look forward to comments from the readers of my blog.

Regards,
Ankur.